EPA-815-R-00-012
DRAFT
TECHNOLOGIES AND COSTS FOR
REMOVAL OF ARSENIC FROM DRINKING WATER
TARGETING AND ANALYSIS BRANCH
STANDARDS AND RISK MANAGEMENT DIVISION
OFFICE OF GROUND WATER AND DRINKING WATER
UNITED STATES ENVIRONMENTAL PROTECTION AGENCY
WASHINGTON, D.C.
NOVEMBER 1999
This report is a draft, issued in support of a proposed National Primary
Drinking Water Regulation for Arsenic. It is intended for public comment and
does not represent final agency policy. EPA expects to issue a final version of
this report in 2001, reflecting corrections due to public comment on the
proposed rule and supporting documents.
INTERNATIONAL CONSULTANTS, INC.
4134 Linden Avenue
Dayton, Ohio 45432
MALCOLM PIRNIE, INC.
432 North 44th Street, Suite 400
Phoenix, Arizona 85008
Under Contract with the USEPA No. 68-C6-0039
Delivery Order 13
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ACKNOWLEDGEMENTS
This document was prepared by the United States Environmental Protection Agency, Office
of Ground Water and Drinking Water under the guidance of the Standards and Risk Management
Division, Targeting and Analysis Branch. The Task Order Project Officer was Mr. Amit Kapadia.
Teclinical consultants played a significant role in the preparation of this document. The
primary technical consultant was International Consultants, Inc. with significant subcontract
assistance provided by Malcolm Pirnie, Inc. The Project Manager was Ron Braun of International
Consultants. The International Consultants technical support team was lead by Chris Hill, Technical
Project Leader, and included Helen Owens, Dr. Dave Jorgenson, Roger Azar and Paul Harvey. The
Malcolm Pirnie technical support team was lead by Zaid Chowdhury with project support from Anne
Jack, Amparo Flores and Spyros Papadimas. Additional technical assistance and consultation were
provided by: Gary Amy, University of Colorado - Boulder; Joe Chwirka, CH2MHD1 - Albuquerque;
Dennis Clifford, University of Houston; Marc Edwards, Virginia Polytechnic Institute; Janet Hering,
California Institute of Technology; Jeff Robinson, Indiana-American Water Company; and Mark
Weisner, Rice University.
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TABLE OF CONTENTS
EXECUTIVE SUMMARY ES-1
Background ~ ES-1
Arsenic Properties and Removal Technologies ES-2
Development of Design Criteria and Treatment Costs ES-4
Residuals Handling and Disposal Alternatives ES-8
Point-of-Entry and Point-of-Use Treatment Options ES-10
Regjonalization ES-11
1.0 INTRODUCTION 1-1
1.1 Overview 1-1
1.2 Statutory Requirements 1-2
1.3 Document Organization 1-3
2.0 ARSENIC REMOVAL TECHNOLOGIES 2-1
2.1 Introduction 2-1
2.2 Precipitative Processes 2-1
2.2.1 Coagulation/Filtration 2-1
2.2.2 Iron/Manganese Oxidation 2-6
2.2.3 Coagulation Assisted Microfiltration 2-7
2.2.4 Enhanced Coagulation 2-7
2.2.5 Lime Softening 2-8
2.3 Adsorptive Processes 2-12
2.3.1 Activated Alumina 2-12
2.3.2 Iron Oxide Coated Sand 2-18
2.4 Ion Exchange 2-20
2.4.1 Introduction 2-20
2.4.2 Effect of pH 2-21
2.4.3 Effect of Competing Ions 2-21
2.4.4 Resin Type 2-22
2.4.5 Process Configuration 2-23
2.4.6 Secondary Effects 2-24
2.4.7 Resin Fouling 2-25
2.4.8 Regeneration 2-25
2.4.9 Regenerant Reuse and Treatment 2-26
2.4.10 EBCT 2-27
2.4.11 Typical Design Parameters 2-27
2.5 Membrane Processes 2-28
2.5.1 Introduction 2-28
2.5.2 Important Factors for Membrane Performance 2-29
U.S. Environmental Protection Agency
Kegion5,Ubrary(pL-i2J)
c"isnxo%7'i2u>fi-
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2.5.3 Arsenic Removal with Membrane Processes 2-31
2.5.4 Microfiltration 2-31
2.5.5 Ultrafiltration 2-32
2.5.6 Nanofiltration 2-34
2.5.7 Reverse Osmosis 2-37
2.5.8 Electrodialysis Reversal 2-40
2.6 Alternative Technologies ". 2-43
2.6.1 Oxidation Filtration 2-43
2.6.2 Sulfur-Modified Iron 2-44
2.6.3 Granular Ferric Hydroxide 2-45
2.6.4 Iron Filings 2-47
2.6.5 Photo-Oxidation 2-47
3.0 TECHNOLOGY COSTS 3-1
3.1 Introduction 3-1
3.2 Basis for Cost Estimates 3-1
3.2.1 Cost Modeling 3-1
3.2.2 Technology Design Panel Recommendations 3-2
3.2.3 Implementing TDP Recommended Costing Upgrades 3-7
3.2.3.1 VSS Model 3-7
3.2.3.2 Water Model 3-8
3.2.3.3 WAV Cost Model 3-8
3.2.4 Cost Indices and Unit Costs 3-9
3.2.5 Re-Basing Bureau of Labor Statistics Cost Indices 3-11
3.2.6 Flows Used in the Development of Costs 3-12
3.3 Additional Capital Costs 3-14
3.4 Costs for Multiple Removal Percentages 3-18
3.4.1 Removal and Accessory Costs 3-18
3.4.2 Use of Blending in Cost Estimates 3-19
3.5 Pre-oxidation Processes 3-20
3.5.1 Potassium Permanganate 3-20
3.5.2 Chlorination 3-28
3.6 Precipitative Processes 3-35
3.6.1 Coagulation/Filtration 3-35
3.6.2 Enhanced Coagulation 3-39
3.6.3 Direct Filtration 3-42
3.6.4 Coagulation Assisted Microfiltration 3-46
3.6.5 Lime Softening 3-50
3.6.6 Enhanced Lime Softening 3-53
3.7 Adsorption Processes 3-56
3.7.1 Activated Alumina 3-56
3.8 Ion Exchange Processes 3-72
3.8.1 Anion Exchange 3-72
3.9 Separation Processes 3-87
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3.9.1 Microfiltration 3-87
3.9.2 Ultrafiltration 3-87
3.9.3 Nanofiltration 3-91
3.9.4 Reverse Osmosis 3-95
3.10 Greensand Filtration 3-98
3.11 Comparison of Costs 3-101
3.11.1 Capital Cost Comparison - 3-101
3.11.2 O&M Cost Comparison 3-102
4.0 RESIDUALS HANDLING AND DISPOSAL ALTERNATIVES 4-1
4.1 Introduction 4-1
4.1.1 Factors Affecting Residuals Handling and Disposal Costs 4-1
4.1.2 Methods for Estimating Residuals Handling and Disposal Costs 4-2
4.2 Residuals Handling Options 4-2
4.2.1 Gravity Thickening 4-2
4.2.2 Mechanical Dewatering 4-3
4.2.3 Evaporation Ponds and Drying Beds 4-4
4.2.4 Storage Lagoons 4-4
4.3 Disposal Alternatives 4-5
4.3.1 Direct Discharge 4-5
4.3.2 Indirect Discharge 4-6
4.3.3 Dewatered Sludge Land Application 4-8
4.3.4 Sanitary Landfill Disposal 4-8
4.3.5 Hazardous Waste Landfill Disposal 4-9
4.4 Residuals Characteristics 4-10
4.4.1 Coagulation/Filtration 4-10
4.4.2 Enhanced Coagulation 4-12
4.4.3 Direct Filtration 4-14
4.4.4 Coagulation Assisted Microfiltration 4-16
4.4.5 Lime Softening 4-17
4.4.6 Enhanced Lime Softening 4-18
4.4.7 Ion Exchange 4-20
4.4.8 Activated Alumina 4-22
4.4.9 Microfiltration 4-24
4.4.10 Ultrafiltration 4-25
4.4.11 Nanofiltration 4-26
4.4.12 Reverse Osmosis 4-28
4.5 Summary 4-30
4.6 Residuals Handling and Disposal Costs 4-33
5.0 POINT-OF-ENTRY/POINT-OF-USE TREATMENT OPTIONS 5-1
5.1 Introduction 5-1
5.2 Variables Affecting Removal Efficiency 5-2
5.2.1 Speciation 5-2
iii
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5.2.2 pH 5-2
5.2.3 Co-occurrence 5-3
5.3 POE/POU Device Case Studies 5-3
5.3.1 Case Study 1: Fairbanks, Alaska and Eugene, Oregon 5-5
5.3.2 Case Study 2: San Ysidro, New Mexico 5-5
5.4 Reverse Osmosis 5-6
5.4.1 Cost Estimates - 5-7
5.5 Ion Exchange 5-9
5.5.1 Cost Estimates 5-9
5.6 Activated Alumina 5-11
5.6.1 Cost Estimates 5-11
6.0 REGIONALIZATION 6-1
6.1 Background 6-1
6.2 Cost Estimates 6-2
7.0 REFERENCES 7-1
APPENDIX A
APPENDIX B
APPENDIX C
APPENDIX D
APPENDIX E
APPENDIX F
APPENDIX G
APPENDIX H
VERY SMALL SYSTEMS CAPITAL COST BREAKDOWN
SUMMARIES
WATER MODEL CAPITAL COST BREAKDOWN SUMMARIES
WAV COST MODEL CAPITAL COST BREAKDOWN
SUMMARIES
COST EQUATIONS AND CURVE FITS FOR REMOVAL AND
ACCESSORY COSTS
ADDITIONAL CAPITAL COSTS
BASIS FOR REVISED ANION EXCHANGEE COSTS
BASIS FOR REVISED ACTIVATED ALUMINA COSTS
REGIONALIZATION COST
IV
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LIST OF TABLES
ES-1 Flows Used in the Cost Estimation Process ES-5
ES-2 Design Criteria and Key Assumptions ES-6
2-1 Typical IX Resins for Arsenic Removal : 2-24
2-2 Typical Operating Parameters and Options for IX 2-27
2-3 Typical Pressure Ranges for Membrane Processes 2-28
2-4 Typical Recovery for Membrane Processes 2-30
2-5 As(V) and As(III) Removal by UF Membranes 2-32
2-6 Arsenic Removal by UF at Pilot-Scale 2-33
2-7 AsO/) and As(III) Removal by NF Membranes 2-35
2-8 Arsenic Removal with NF at Pilot-Scale 2-36
2-9 Summary of Arsenic Removal with RO 2-38
2-10 Arsenic Removal with RO at Bench-Scale 2-39
2-11 Arsenic Removal with RO at Pilot-Scale 2-40
2-12 Influent Water Quality for San Ysidro EDR Study 2-42
2-13 Raw Water Quality for Bluewater EDR Study 2-43
2-14 Adsorption Tests on GFH 2-46
3-1 TOP Capital Cost Factors 3-3
3-2 VSS Capital Cost Breakdown for Membrane Processes 3-4
3-3 Water Model Capital Cost Breakdown for Package Conventional Treatment 3-5
3-4 Wat;r Model Capital Cost Breakdown by Percentage for Package Conventional
Treatment 3-5
3-5 W/W Cost Model Capital Cost Breakdown for Sedimentation Basins 3-6
3-6 W/W Cost Model Capital Cost Breakdown by Percentage for Sedimentation Basins . 3-6
3-7 Cost Indices Used in the Water and WAV Cost Models 3-9
3-8 Unit and General Cost Assumptions 3-9
3-9 Chemical Costs 3-10
3-10 Amortization Factors 3-11
3-11 Bureau of Labor Statistics Re-base Information 3-12
3-12 Flows Used in the Cost Estimation Process 3-13
3-13 Permitting Scenarios 3-16
3-14 Treatment Technology Maximum Achievable Removal Percentages 3-18
3-15 Regeneration Frequency vs. Influent Arsenic Concentration for Activated Alumina . 3-56
3-16 Influent pH vs. Regeneration for Activated Alumina 3-84
3-17 Number of DC Beds Included in Cost Estimates 3-80
4-1 Summary of Residuals Characteristics 4-31
4-2 Summary of Arsenic Residuals Handling and Disposal Options 4-32
5-1 Source Water Summary - Point-of-Use Case Studies 5-4
5-2 Observed Arsenic Removal by Technology for POE and POU Units 5-4
6-1 Regionalization Cost Estimates 6-3
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LIST OF FIGURES
2-1 Pressure Driven Membrane Process Classification 2-29
3-1 Preoxidation -1.5 mg/L Permanganate Capital Costs 3-22
3-2 Preoxidation -1.5 mg/L Permanganate O&M Costs " 3-23
3-3 Preoxidation - 3.0 mg/L Permanganate Capital Costs 3-24
3-4 Preoxidation - 3.0 mg/L Permanganate O&M Costs 3-25
3-5 Preoxidation - 5.0 mg/L Permanganate Capital Costs 3-26
3-6 Preoxidation - 5.0 mg/L Permanganate O&M Costs 3-27
3-7 Preoxidation -1.5 mg/L Chlorine Capital Costs 3-29
3-8 Preoxidation -1.5 mg/L Chlorine O&M Costs 3-30
3-9 Preoxidation - 3.0 mg/L Chlorine Capital Costs 3-31
3-10 Preoxidation - 3.0 mg/L Chlorine O&M Costs 3-32
3-11 Preoxidation - 5.0 mg/L Chlorine Capital Costs 3-33
3-12 Preoxidation - 5.0 mg/L Chlorine O&M Costs 3-34
3-13 Coagulation/Filtration Capital Cost 3-37
3-14 Coagulation/Filtration O&M Cost 3-38
3-15 Enhanced Coagulation/Filtration Capital Cost 3-40
3-16 Enhanced Coagulation/Filtration O&M Cost 3-41
3-17 Direct Filtration Capital Cost 3-44
3-18 Direct Filtration O&M Cost 3-45
3-19 Coagulation Assisted Microfiltration Capital Cost 3-48
3-20 Coagulation Assisted Microfiltration O&M Cost 3-49
3-21 Lime Softening Capital Cost 3-51
3-22 Lime Softening O&M Cost 3-52
3-23 Enhanced Lime Softening Capital Cost 3-54
3-24 Enhanced Lime Softening O&M Cost 3-55
3-25 Activated Alumina, No Regeneration Capital Costs 3-60
3-26 Activated Alumina, No Regeneration O&M Costs 3-61
3-27 Activated Alumina Capital Cost 3-62
3-28 Activated Alumina O&M Cost - 500 BV 3-63
3-29 Activated Alumina O&M Cost - 2000 BV 3-64
3-30 Activated Alumina O&M Cost - 3000 BV 3-65
3-31 Activated Alumina O&M Cost - 5000 BV 3-66
3-32 Activated Alumina O&M Cost - 7000 BV 3-67
3-33 Activated Alumina O&M Cost -10000 BV 3-68
3-34 Activated Alumina O&M Cost -16500 BV 3-69
3-35 Activated Alumina O&M Cost - 25000 BV 3-70
3-36 Activated Alumina O&M Cost - 50000 BV 3-71
3-37 Bed Volumes to Arsenic Breakthrough as a Function of Sulfate Concentration 3-73
3-38 Ion Exchange Regeneration Frequency, 50 ppb Influent Arsenic 3-74
3-39 Ion Exchange Regeneration Frequency, 30 ppb Influent Arsenic 3-75
vi
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3-40 Ion Exchange Regeneration Frequency, 20 ppb Influent Arsenic 3-76
3-41 Ion Exchange Regeneration Frequency, 10 ppb Influent Arsenic 3-77
3-42 Ion Exchange Capital Cost 3-81
3-43 Ion Exchange O&M Cost - 300 BV 3-82
3-44 Ion Exchange O&M Cost - 500 BV 3-83
3-45 Ion Exchange O&M Cost - 700 BV 3-84
3-46 Ion Exchange O&M Cost -1500 BV , 3-85
3-47 lor. Exchange O&M Cost - 2500 BV 3-86
3-48 Ultrafiltration Capital Cost 3-89
3-49 Ultrafiltration O&M Cost 3-90
3-50 Nanofiltration Capital Cost 3-93
3-51 Nanofiltration O&M Cost 3-94
3-52 Reverse Osmosis Capital Cost 3-96
3-53 Reverse Osmosis O&M Cost 3-97
3-54 Greensand Filtration Capital Cost 3-99
3-55 Greensand Filtration O&M Cost 3-100
3-56 Comparison of Capital Cost Estimates - Coagulation/Filtration 3-104
3-57 Comparison of O&M Cost Estimates - Coagulation/Filtration 3-105
3-58 Comparison of Capital Cost Estimates - Direct Filtration 3-106
3-59 Comparison of O&M Cost Estimates - Direct Filtration 3-107
3-60 Comparison of Capital Cost Estimates - Lime Softening 3-108
3-61 Comparison of O&M Cost Estimates - Lime Softening 3-109
3-62 Comparison of Capital Cost Estimates - Activated Alumina 3-110
3-63. Comparison of O&M Cost Estimates - Activated Alumina 3-111
3-64 Comparison of Capital Cost Estimates - Ion Exchange 3-112
3-65 Comparison of O&M Cost Estimates - Ion Exchange 3-113
3-66 Comparison of Capital Cost Estimates - Ultrafiltration 3-114
3-67 Comparison of O&M Cost Estimates - Ultrafiltration 3-115
3-68 Comparison of Capital Cost Estimates - Nanofiltration 3-116
3-69 Comparison of O&M Cost Estimates - Nanofiltration 3-117
3-70 Comparison of Capital Cost Estimates - Reverse Osmosis 3-118
3-71 Comparison of O&M Cost Estimates - Reverse Osmosis 3-119
4-1 Mechanical Dewatering and Non-hazardous Landfill, Coagulation-Assisted
Microfiltration, Disposal Capital Costs 4-35
4-2 Mechanical Dewatering and Non-hazardous Landfill, Coagulation-Assisted
Microfiltration, Disposal O&M Costs 4-36
4-3 Non-Mechanical Dewatering and Non-hazardous Landfill, Coagulation-Assisted
Microfiltration, Disposal Capital Costs 4-37
4-4 Non-Mechanical Dewatering and Non-hazardous Landfill, Coagulation-Assisted
Microfiltration, Disposal O&M Costs 4-38
4-5 POTW Discharge - 500', Ion Exchange, Disposal Capital Costs 4-39
4-6 POTW Discharge - 500', Ion Exchange, Disposal O&M Costs 4-40
4-7 Evaporation Pond and Non-hazardous Landfill, Ion Exchange,
Disposal Capital Costs 4-41
Vll
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4-8 Evaporation Pond and Non-hazardous Landfill, Ion Exchange;,
Disposal O&M Costs 4-42
4-9 Chemical Precipitation and Non-hazardous Landfill, Ion Exchange,
Disposal Capital Costs 4-43
4-10 Chemical Precipitation and Non-hazardous Landfill, Ion Exchange,
Disposal O&M Costs 4-44
4-11 Non-hazardous Landfill, Activated Alumina (3000 BV), Disposal O&M Costs 4-45
4-12 Non-hazardous Landfill, Activated Alumina (7000 BV), Disposal O&M Costs 4-46
4-13 Non-hazardous Landfill, Activated Alumina (16500 BV), Disposal O&M Costs ... 4-47
4-14 POTW Discharge (500') and Non-hazardous Landfill, Activated Alumina (3000 BV),
Disposal Capital Costs 4-48
4-15 POTW Discharge (5001) and Non-hazardous Landfill, Activated Alumina (3000 BV),
Disposal O&M Costs 4-49
4-16 POTW Discharge (500') and Non-hazardous Landfill, Activated Alumina (7000 B V),
Disposal Capital Costs 4-50
4-17 POTW Discharge (500') and Non-hazardous Landfill, Activated Alumina (7000 BV),
Disposal O&M Costs 4-51
4-18 POTW Discharge (500') and Non-hazardous Landfill, Activated Alumina (16500 BV),
Disposal Capital Costs 4-52
4-19 POTW Discharge (500') and Non-hazardous Landfill, Activated Alumina (16500 BV),
Disposal O&M Costs 4-53
4-20 Direct Discharge - 500' of Pipe, Reverse Osmosis, Disposal Capital Costs 4-54
4-21 Direct Discharge - 500' of Pipe, Reverse Osmosis, Disposal O&M Costs 4-55
4-22 POTW Discharge - 500', Reverse Osmosis, Disposal Capital Costs 4-56
4-23 POTW Discharge - 500', Reverse Osmosis, Disposal O&M Costs 4-57
4-24 Chemical Precipitation and Non-heizardous Landfill, Reverse Osmosis,
Disposal Capital Costs 4-58
4-25 Chemical Precipitation and Non-hazardous Landfill, Reverse Osmosis,
Disposal O&M Costs 4-59
4-26 POTW Discharge - 500', Greensand Filtration, Disposal Capital Costs 4-60
4-27 POTW Discharge - 500', Greensand Filtration, Disposal O&M Costs 4-61
5-1 POE and POU Total Costs, Reverse Osmosis 5-8
5-2 POE and POU Total Costs, Ion Exchange 5-10
5-3 POE and POU Total Costs, Activated Alumina 5-12
Vlll
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LIST OF ACRONYMS
AA activated alumina
AWWA American Water Works Association
AWWARF American Water Works Association Research Foundation
BLS Bureau of Labor Statistics
BV bed volume
C/F coagulation/filtration
CFR Code of Federal Regulations
D/DBP Disinfectant/Disinfection By-Product
DBF disinfection by-product
DD direct discharge
DMAA dimethyl arsenic acid
DOC dissolved organic carbon
DWRD Drinking Water Research Division
EBCT empty bed contact time
ED electrodialysis
EDR electrodialysis reversal
ENR Engineering News Record
EP evaporation ponds and drying beds
EPA United States Environmental Protection Agency
Fe/Mn iron/manganese
ft feet
GAC granular activated carbon
GFH granular ferric hydroxide
gpd gallons per day
IX
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gpm
GT
HD
HDPE
ID
IOCS
IX
kgal
kgpd
kWh
LA
Ib/acre
LS
MCL
MCLG
MD
MDL
MF
mg/kg
mg/L
MGD or mgd
MMAA
MWCO
MWDSC
NF
NOM
NIPDWR
gallons per minute
gravity thickening
hazardous waste landfill disposal
high-density polyethylene
indirect discharge
iron oxide coated sand
ion exchange
thousand gallons
thousand gallons per day
kilowatt hour
land application
pounds per acre
lime softening
Maximum Contaminant Level
Maximum Contaminant Level Goal
mechanical dewatering
minimum detection limit
micro filtration
milligram per kilogram
milligrams per liter
million gallons per day
monomethyl arsenic acid
molecular weight cut-off
Metropolitan Water District of Southern California
nanofiltration
natural organic matter
National Interim Primary Drinking Water Regulation
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NPDES
NPDWR
O&M
OGWDW
POE
POTW
POU
ppb
ppm
PPI
psi
psig
PVC
RCRA
RO
scf
SD
SDWA
sf
SL
SMI
SOC
sqft
TBLL
TCLP
TDS
TOC
National Pollutant Discharge Elimination System
National Primary Drinking Water Regulation
operations and maintenance
Office of Ground Water and Drinking Water
point-of-entry
public-owned treatment works
point-of-use
parts per billion
parts per million
Producer Price Index (for Finished Goods)
pounds per square inch
pounds per square inch gauge
polyvinyl chloride
Resource Conservation and Recovery Act
reverse osmosis
standard cubic feet
sanitary landfill disposal
Safe Drinking Water Act
square feet
storage lagoons
Sulfur-Modified Iron
synthetic organic compound
square feet
Technically Based Local Limits
Toxicity Characteristic Leaching Procedure
total dissolved solids
total organic carbon
XI
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TSS total suspended solids
TWG Technologies Working Group
UF ultrafiltration
UV254 ultraviolet 254
WET Whole Effluent Toxicity
wk week
yr year
micrograms per liter
Xll
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EXECUTIVE SUMMARY
BACKGROUND
In ] 976 EPA issued a National Interim Primary Drinking Water Regulation (NIPDWR) for
arsenic at :50 parts per billion (ppb or //g/L). Under the 1986 amendments to the Safe Drinking
Water Act (SDWA), Congress directed EPA to publish Maximum Contaminant Level Goals
(MCLGs) and promulgate National Primary Drinking Water Regulations (NPDWRs) for 83
contaminants, including arsenic. As EPA missed the statutory deadline for promulgating an arsenic
regulation, a citizens' group filed suit to compel EPA to do so; EPA entered into a consent decree
to issue the regulation. The EPA Office of Ground Water and Drinking Water (OGWDW) held
internal workgroup meetings throughout 1994, addressing risk assessment, treatment, analytical
methods, arsenic occurrence, exposure, costs, implementation issues, and regulatory options before
deciding m early 1995 to defer the regulation to better characterize health effects and treatment
technology. When Congress reauthorized the SDWA on the August 6,1996, section 1412(b)(12)(A)
was added. This addition specifies in part, that EPA propose a NPDWR for arsenic by January 1,
2000 and issue a final regulation by January 1, 2001.
The purpose of this document is to characterize the ability of arsenic removal technologies
and to estimate costs for technologies that can be used by utilities to meet regulatory standards. This
document was originally published in 1993 as Treatment and Occurrence of Arsenic in Potable
Water Supplies (Malcolm Pirnie, 1993a). Design criteria from the 1993 document have been re-
evaluated and modified in accordance with the most recent research and input from a panel of
experts. The design criteria established were used to develop treatment costs which will be used by
EPA to det<;rmine national costs for various arsenic regulatory scenarios.
Costs were developed using the WAV Cost Model (Culp/Wesner/Culp, 1994), the Water
Model (Culp/Wesner/Culp, 1984), and the Very Small Systems Best Available Technology Cost
Document (Malcolm Pirnie, 1993b). For some technologies (e.g., membranes), published data of
operating plants were used to estimate costs, as the models were judged inadequate or out of date.
Where appropriate, vendors and equipment manufacturers were contacted to assess the accuracy of
the cost models, and, when necessary, costs were modified to reflect the input from these sources.
ES-1
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ARSENIC PROPERTIES AND REMOVAL TECHNOLOGIES
Arsenic occurs in two primary forms; organic and inorganic, i^rganic species of arsenic are
predominantly found in foodstuffs, such as shellfish, and include such forms as monomethyl arsenic
acid (MMAA), dimethyl arsenic acid (DMAA), and arseno-sugars. Inorganic arsenic occurs in two
valence states, arsenite (As III) and arsenal e (As V). As(III) species consist primarily of arsenious
acid (H3AsO3) in natural waters. As(V) sprcies consist primarily of H2AsO4' and HAsO42" in natural
waters (Clifford and Lin, 1995). Most natural waters contain the inorganic forms of arsenic.
Moreover, natural groundwaters contain the more toxic (among the inorganic species) form As(III)
as reducing conditions prevail. In natural surface waters, however, As(V) is the dominant species.
Arsenite is removed less efficiently because: it predominantly occurs in the uncharged (H3AsO3) state
in source waters with a pH of less than 9.0. The dominant arsenate forms are anionic species,
H2AsO42- and HAsO4'.
Arsenic removal is dependent upon the ionic form present and water chemistry. As a result,
identification of the ionic form is necessary for selection and design of an arsenic removal process.
All technologies discussed in this document remove arsenate more effectively than arsenite.
Therefore, if arsenite is the predominant species present, oxidation to arsenate may be required to
achieve the desired removal.
Source water pH plays a significant role in determining the removal efficiency of a particular
technology. Most processes are relatively unaffected by pH in the range of 6.5 to 9.0. However,
activated alumina studies have shown the optimum pH for arsenic removal to be between 5.5 and
6.0. Reverse osmosis processes may requite pH adjustment to prevent precipitation of salts on the
membrane surface.
Co-occurrence of inorganic contaminants, such as sulfate and silica, as well as suspended
solids, can cause interference with arsenic removal. Sulfate is preferentially adsorbed over arsenic
by ion exchange processes. This preference can result in another phenomenon known as peaking,
which occurs when arsenic is displaced on 1he resins by the sulfate causing effluent concentrations
in excess of the influent levels.
ES-2
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This document evaluates arsenic removal technologies for drinking water and the costs
associated with those technologies. Specifically, the following treatment processes are discussed:
• Precipitative processes, including coagulation/filtration (C/F), direct filtration,
coagulation assisted microfiltration, enhanced coagulation, lime softening (LS), and
enhanced lime softening;
• Adsorption processes, including activated alumina (AA), and iron oxide coated sand
(IOCS);
• Ion exchange (IX) processes, specifically anion exchange;
• Membrane filtration, including microfiltration (MF), ultrafiltration (UF),
nanofiltration (NF), reverse osmosis (RO), and electrodialysis reversal (EDR);
• Alternative treatment processes, including biological processes, granular ferric
hydroxide, sulfur-modified iron and iron filings, greensand filtration, and photo-
oxidation; and
• Point-of-entry (POE) and point-of-use (POU) devices.
Many of these processes were evaluated to develop cost curves for the technologies. The
above list also includes some experimental technologies which could not be costed out at this time.
Discussions of each of these technologies are included in this document for future considerations of
these proceisses for arsenic removal:
• Iron oxide coated sand;
• Iron filings and sulfur-modified iron;
• Granular ferric hydroxide; and
• Photo-oxidation.
ES-3
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DEVELOPMENT OF DESIGN CRITERIA AND TREATMENT COSTS
Three cost models were used in cost development: the Very Small Systems Best Available
Technology Cost Document (Malcolm Pirnie, 1993), hereafter referred to as the VSS model; the
Water Model (Culp/Wesner/Culp, 1984); and the WAV Cost Model (CulpAVesner/Culp, 1994).
Curve fitting analysis was conducted on 'the modeled cost estimates including the utilization of
transition flow regions to provide better estimates within the breakpoints between models. The
following flow ranges have been established for each model and transition flow region:
VSS - 0.015 to 0.100 mgd
Transition 1 - 0.100 to 0.270 mgd
Water Model - 0.27 to 1.00 mgd
Transition 2 - 1 to 10 mgd
WAV Cost Model - 10 to 200 mgd
Flow categories were developed to provide adequate characterization of costs across each of
the flow regions presented above. A minimum of four data points were generated for each of the
flow regions, with the exception of the transition regions, where cost estimates are based upon a
linear regressions between the last data point of the previous region and the first data point of the
following region. Table ES-1 presents the design and average flows, .and cost models used in this
process.
The arsenic species present can greatly affect the removal efficiisncy of the selected treatment
process. Pre-oxidation may be necessary to convert arsenite to arsenate. This document presents
two pre-oxidation alternatives; chlorination and potassium permanganate feed. Costs are presented
for dosages of 1.5, 3.0 and 5.0 mg/L for both pre-oxidation teclinologies. Other oxidation
technologies, such as hydrogen peroxide feed and ozonation, may dso be effective, but are not
considered as typical oxidation processes for this contaminant.
Design criteria for the technologies selected for the development of cost curves are shown
in Table ES-2. These design criteria were developed by consulting various engineers and experts
convened by the American Water Works Association (AWWA) in 1994 (Frey, et al., 1997).
ES-4
-------
Table ES-1
Flows Used in the Cost Estimation Process
Design Flow (mgd)
0.010
0.024
0.087
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0.27
0.45
0.65
0.83
1,0:
1.8
4.8
10
11
18
26
51
210
430
Average Flow (mgd)
0.0031
0.0056
0.024
0.031
0.086
0.14
0.23
0.30
.. ,. 0.36 :
0.7
2.1
4.5
5
8.8
13
27
120
270
Cost Model
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Water
Water
Water
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W/W Cost
WAV Cost
WAV Cost
WAV Cost
WAV Cost
WAV Cost
WAV Cost
WAV Cost
WAV Cost
Shaded rows represent data used in the estimation of costs with the transition regions.
ES-5
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Expansion during regeneration = 50%
Regeneration cycle = 10 minutes backwashing at
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ES-7
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Capital costs are presented in April 1998 dollars. Appropriate Engineering News Record
(ENR) and Bureau of Labor Statistics (BLS) cost indices were used for cost computation. The
Producer's Price Index for Finished Goods were used adjusting operations and maintenance (O&M)
cost estimates.
Capital and O&M cost curves and equations are presented for each technology discussed in
Chapter 3. Capital costs are expressed as total cost (M-$); O&M estimates are expressed in dollars
per year. For the ion exchange and activated alumina processes, a range of O&M costs have been
cited since costs will vary with the number of bed volumes treated between regeneration, which is
a function of several water quality parameters, such as ambient sulfate level, and initial and target
arsenic concentrations.
Capital and O&M cost estimates were compared with actual data presented in Evaluation of
Full-Scale Treatment Technologies at Small Drinking Water Systems (ICF and ISSI, 1998). It was
found that the estimates presented in this document are reasonable. Capital cost estimates were
routinely conservative, but followed the general trends seen in actual data. O&M estimates typically
represented an approximate average of the real world costs. Actual data was not available for all
technologies, and comparisons are not presented for some of the technologies discussed.
RESIDUALS HANDLING AND DISPOSAL ALTERNATIVES
Each of the treatment technologies presented in this document will produce residuals, either
solid or liquid streams, containing elevated levels of arsenic. It is important to address residuals
characteristics when selecting an arsenic removal technology. Handling and disposal costs can be
significant, and if a waste stream happens to be hazardous the implications are even greater. This
document evaluates typical characteristics of residuals produced by each of the treatment
technologies presented, and discusses appropriate handling and disposal methods. Specifically, the
following handling and disposal methods are discussed:
• Residuals handling
Gravity thickening;
Mechanical dewatering, including centrifuges and filter presses;
Non-mechanical dewatering, such as evaporation ponds and storage lagoons;
ES-8
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• Disposal alternatives
Direct discharge to receiving water;
Discharge to sanitary sewer for treatment at a wastewater treatment plant;
Land application;
Sanitary landfill disposal; and
Hazardous landfill disposal.
There are a number of factors which can influence residuals handling and disposal costs. The
primary factor affecting capital cost is the size of the water system, i.e., population and water needs.
All other costs are directly proportional to these factors. The amount of waste generated plays a
significant role in determining the handling and disposal method to be utilized. Many handling
methods are impractical for large water systems because of land requirements. However, some
handling methods require expensive process equipment which may make them more suitable to large
water systems. Similarly, waste disposal methods requiring large capital investments may make
them impractical for small water systems.
Many handling and disposal methods require extensive oversight which can be a burden on
small water systems. Generally, labor intensive technologies are more suitable to large water
systems. Transportation can also play a significant role in determining appropriate handling and
disposal options. If off-site disposal requires extensive transportation, alternative disposal methods
should be evaluated. Complex handling and disposal methods usually require more maintenance.
When evaluating handling and disposal methods, it is generally best to select that option which will
require the least amount of oversight and maintenance.
Residuals handling and disposal costs can be difficult to estimate. There are a number of
factors which affect capital and O&M costs, and disposal costs can be largely regional. EPA has
published two manuals for estimating residuals handling and disposal costs; Small Water System
Byproducts Treatment and Disposal Cost Document (DPRA, 1993 a), and Water System Byproducts
Treatment and Disposal Cost Document (DPRA, 1993b). Both present a variety of handling and
disposal options, applications and limitations of those technologies, and capital and O&M cost
equations. Residuals handling and disposal costs are not included in this document. The references
listed above can be used to generate such costs.
ES-9
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POINT-OF-ENTRY AND POINT-OF-USE TREATMENT OPTIONS
Centralized treatment is not always a feasible treatment option, for example in areas where
each home has a private well or centralized treatment is cost prohibitive. In these instances, point-
of-entry (POE) and point-of-use (POU) treatment options may be acceptable treatment alternatives.
POE and POU systems offer ease of installation, simplify operation and maintenance, and generally
have lower capital costs. These systems may also reduce engineering, legal and other fees typically
associated with centralized treatment options. Use of POE and POU systems does not reduce the
need for a well-maintained water distribution system. In fact, increased monitoring may be
necessary to ensure that the treatment units are operating properly.
Home water treatment can consist of either whole-house or single faucet treatment. Whole-
house, or POE treatment is necessary when exposure to the contaminant by modes other than
consumption is a concern. POU treatment is preferred when treated water is needed only for
drinking and cooking purposes. POU treatment usually involves single-tap treatment.
Section 1412(b)(4)(E) of the 1996 Safe Drinking Water Act (SDWA) Amendments requires
the EPA to issue a list of technologies that achieve compliance with Maximum Contaminant Levels
(MCLs) established under the act. This list must contain technologies for each NPDWR and for each
of the small public water systems categories listed below:
• Population of more than 50, but less than 500;
• Population of more than 500, but less than 3,300; and.
• Population of more than 3,3 00, but less than 10,000.
The SDWA identifies POE and POU treatment units as potentially affordable technologies,
but stipulates that POE and POU treatment systems "shall be owned, controlled and maintained by
the public water system, or by a person under contract with the public water system to ensure proper
operation and compliance with the maximum contaminant level or treatment technique and equipped
with mechanical warnings to ensure that customers are automatically notified of operational
problems."
ES-10
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Research has shown that POE and POU devices can be effective means of removing arsenic
from potable water. Water systems with high influent arsenic concentrations, i.e., greater than 1
mg/L, may have difficulty meeting MCLs much lower than the 10 to 20 f-tgfL level. As a result,
influent arsenic concentration and other source water characteristics must be considered when
evaluating POE and POU devices for arsenic removal. Reverse osmosis, activated alumina and ion
exchange are three treatment techniques that have been evaluated and shown to be effective. This
document looks at the removals achieved by each of these three treatment techniques, and presents
total costs for each treatment option.
REGIONALIZATION
Regionalization involves purchasing and transferring water from one community or water
source to another. In effect, regionalization expands the region served by a water distribution
system. ITiere are a number of factors which can influence the decision to implement
regionalizalion, including water availability, water quality, geography and economic factors.
Accordingly, community water systems faced with installation of treatment facilities to
address arsenic contamination issues may opt for regionalization. This document presents the costs
associated with regionalization. Costs can be largely variable, with many site specific
considerations. The estimates presented are for typical installation and do not include costs
associated with site specific construction conditions, such as rugged terrain, severe elevation changes
and land costs.
ES-11
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1.0 INTRODUCTION
1.1 OVERVIEW
Arsenic (As) is a naturally occurring element present in food, water, and air. Known for
centuries to be an effective poison, some animal studies suggest that arsenic may be an essential
nutrient at low concentrations. Non-malignant skin alterations, such as keratosis and hypo- and
hyper-pigmentation, have been linked to arsenic ingestion, and skin cancers have developed in some
patients. Additional studies indicate that arsenic ingestion may result in internal malignancies,
including cancers of the kidney, bladder, liver, lung, and other organs. Vascular system effects have
also been observed, including peripheral vascular disease, which in its most severe form, results in
gangrene or Blackfoot Disease. Other potential effects include neurologic impairment (Lomaquahu
and Smith, 1998).
The primary route of exposure to arsenic for humans is ingestion. Exposure via inhalation
is considered minimal, though there are regions where elevated levels of airborne arsenic occur
periodically (Hering and Chiu, 1998). Arsenic occurs in two primary forms; organic and inorganic.
Organic species of arsenic are predominantly found in foodstuffs, such as shellfish, and include such
forms as monomethyl arsenic acid (MMAA), dimethyl arsenic acid (DMAA), and arseno-sugars.
Inorganic arsenic occurs in two valence states, arsenite (As III) and arsenate (As V). As(III) species
consist primarily of arsenious acid (H3AsO3) in natural waters. As(V) species consist primarily of
H2AsO4" and HAsO42' in natural waters (Clifford and Lin, 1995). Most natural waters contain the
more toxic inorganic forms of arsenic. Natural groundwaters contain predominantle As(III) since
reducing conditions prevail. In natural surface waters As(V) is the dominant species. Arsenic
removal teclinologies for drinking water include:
Precipitative processes, including coagulation/filtration (C/F), direct filtration,
coagulation assisted microfiltration, enhanced coagulation, lime softening (LS), and
enhanced lime softening;
Adsorption processes, including activated alumina (AA), and iron oxide coated sand
(IOCS);
l-l
-------
" Ion exchange (IX) processes, specifically anion exchsmge;
• Membrane filtration, including microfiltration (MF), ultrafiltration (UF),
nanofiltration (NF), reverse osmosis (RO), and electrodialysis reversal (EDR);
• Alternative treatment processes, including biological processes, granular ferric
hydroxide, sulfur-modified iron and iron filings, and greensand filtration; and
• Point-of-entry (POE) and point-of-use (POU) devices.
Many of these processes were evaluated to develop cost curves for the technologies. The
above list also includes some experimental technologies which could not be costed at this time.
Discussions of the following technologies are included in this document for future consideration as
viable processes for arsenic removal:
• Iron oxide coated sand;
• Greensand filtration;
• Iron filings and sulfur-modified iron; and
• Granular ferric hydroxide.
1.2 STATUTORY REQUIREMENTS
In 1976 EPA issued a National Interim Primary Drinking Water Regulation (NIPWDR) for
arsenic at 50 parts per billion (ppb or //g^L). Under the 1986 amendments to the Safe Drinking
Water Act (SDWA), Congress directed EPA to publish Maximum Contaminant Level Goals
(MCLGs) and promulgate National Primary Drinking Water Reflations (NPDWRs) for 83
contaminants, including arsenic. When EPA missed the statutory (deadline for promulgating an
arsenic regulation, a citizens' group filed suit to compel EPA to do so; EPA entered into a consent
decree to issue the regulation. The EPA Office of Ground Water and Drinking Water (OGWDW)
held internal workgroup meetings throughout 1994, addressing risk assessment, treatment, analytical
methods, arsenic occurrence, exposure, cos;ts, implementation issues, and regulatory options before
1-2
-------
deciding in early 1995 to defer the regulation to better characterize health effects and treatment
technology.
With the reauthorization of the SDWA on August 6, 1996, Congress added section
1412(b)(12)(A) to the act. This addition specifies, in part, that EPA propose a NPDWR for arsenic
by January 1,2000 and issue a final regulation by January 1,2001.
1.3 DOCUMENT ORGANIZATION
This document contains the following chapters:
Chapter 1.0 Introduction - Provides an introduction to the arsenic statutory requirements,
and defines technology categories. Also presents the organization of the document.
Chapter 2.0 Arsenic Removal Technologies - Presents discussions on available arsenic
removal technologies, removal efficiencies, factors affecting arsenic removal, and associated
pilot- and full-scale studies.
Chapter 3.0 Technology Costs - Presents capital and O&M costs for each of the removal
technologies in graphical format. This chapter also contains a comparison of the cost
estimates presented in this document to actual capital and O&M costs obtained during an
EPA survey of small water systems.
Chapter 4.0 Residuals Handling and Disposal Alternatives - Presents capital and O&M
cost equations for a variety of residuals handling and disposal alternatives.
Chapter 5.0 Point-of-Entry/Point-of-Use Treatment Options - Evaluates a number of
POE and POU treatment options effective for arsenic removal, as well as presents capital and
O&M costs in graphical form for each of the treatment options.
Chapter 6.0 Regionalization - Presents estimates for regionalization as opposed to
centralized treatment.
Chapter 7.0 References - Lists the literature cited in this document, as well as additional
references which may be of interest to the reader.
1-3
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2.0 ARSENIC REMOVAL TECHNOLOGIES
2.1 INTRODUCTION
Arsenic removal technologies are discussed in this chapter. Some of these technologies are
traditional treatment processes which have been tailored to improve removal of arsenic. Several
treatment techniques discussed here are at the experimental stage with regard to arsenic removal, and
some have not been demonstrated at full scale. Although some processes may be technically
feasible, cost may be prohibitive.
Teclmologies discussed here are grouped into four broad categories: precipitative processes,
adsorption processes, ion exchange processes, and separation (membrane) processes. Each category
is discussed here, with at least one treatment technology described in each category.
2.2 PRECIPITATIVE PROCESSES
2.2.1 Coagulation/Filtration
Coagulation/filtration is a treatment process by which the physical or chemical properties of
dissolved colloidal or suspended matter are altered such that agglomeration is enhanced to an extent
that the resulting particles will settle out of solution by gravity or will be removed by filtration.
Coagulants change surface charge properties of solids to allow agglomeration and/or enmeshment
of particles into a flocculated precipitate. In either case, the final products are larger particles, or
floe, which more readily filter or settle under the influence of gravity.
The
-------
handling equipment, and filter backwash facilities. Settling may not be necessary in situations where
the influent particle concentration is very low. Treatment plants without settling are known as direct
filtration plants.
As(III) removal during coagulation with alum, ferric chloride, and ferric sulfate has been
shown to be less efficient than As(V) under comparable conditions (Hering, et al., 1996; Edwards,
1994; Shen, 1973; Gulledge and O'Conner, 1973; Sorg and Logsdon, 1978). If only As(III) is
present, consideration should be given to oxidation prior to coagulation to convert As(III) to As(V)
species.
Effect of CoaguIanUype
Batch studies were conducted at the University of Illinois to demonstrate the removal of
As(V) by coagulation, sedimentation, and filtration (Gulledge and O'Conner, 1973). Raw water
was spiked to obtain an initial concentration of 0.05 mg/L As(V); alum or ferric sulfate were used
as coagulants at varying dosages. The pH was varied between 5.0 zmd 8.0, which is higher than
the optimum pH range of 5.0 to 7.0 for alum coagulation, but within the optimum pH range for
ferric sulfate coagulation. The results of these studies demonstrate that ferric sulfate coagulation
within the optimum pH range achieved better removals than alum coagulation over a larger
coagulant dosage range. Over 90 percent of As(V) was removed with alum coagulation but only
at dosages greater than 30 mg/L. With ferric sulfate coagulation, over 95 percent of the As(V)
was removed within the pH range of 5.0 to 7.5 for dosages between 10 and 50 mg/L.
Logsdon et al. (1974) showed that at an influent concentration of 0.3 mg/L, removals
ranged from 40 to 60 percent with ferric sulfate coagulation, compared to 5 to 15 percent with
alum coagulation. Higher As(m) removals were achieved in the pH range of 5.0 to 8.5 for ferric
sulfate and 5.0 to 7.0 for alum. When As(HI) was oxidized with 2 mg/L of chlorine, removals
increased for both alum and ferric sulfate within the same pH range, but ferric sulfate still
achieved higher removals. Over 95 percent of the oxidized As(IH) was removed with ferric
sulfate coagulation, and between 83 and 90 percent was removed with alum coagulation.
Scott, et al. (1995) conducted a full-scale study at the Metropolitan Water District of
Southern California (MWDSC) to determine arsenic removals using alum and ferric chloride. The
2-2
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average concentration of arsenic in the source water was 2.1 ug/L. When the source water was
treated with 3 to 10 mg/L of ferric chloride, arsenic removal was 81 to 96 percent. When the source
water was treated with 6,10, or 20 mg/L of alum, arsenic removal was 23 to 71 percent.
McNeill and Edwards (1997a) reported that solubility and stability of the metal hydroxide
floes play an important role hi arsenic removal. When ferric coagulants are added, most of the
ferric ends; up as ferric hydroxide. In alum coagulation, however, a significant portion of the
added aluminum remains as soluble complexes. Because only particulate metal hydroxides can
mediate arsenic removal, alum plants must carefully consider aluminum solubility when arsenic
removal is required. Aluminum complexes can pass through filters and decrease overall arsenic
removal.
Effect of Coagulant Dosage
In general, higher removal efficiencies can be achieved with increased coagulant dosages
(Cheng, et al., 1994; Edwards, 1994; Gulledge and O'Conner, 1973). Hering et al. (1996)
demonstrate in coagulation experiments with ferric chloride at pH 7.0 that both As(III) and As(V)
removal were dependent on coagulant dosage. "Complete" removal of As(V) was observed for
coagulant dosages above 5 mg/L ferric chloride.. "Complete" removal of As(III) was not observed
under the range of conditions examined.
Predictions based on existing data and the use of a diffuse-layer model indicated that As(III)
removals by coagulation were primarily controlled by coagulant dosage, whereas the converse was
true for As(V) (Edwards, 1994). A database compiled by Edwards (1994) containing much
previously published work on arsenic coagulation indicated that, at all dosages greater than 20 mg/L
as ferric chloride or 40 mg/L as alum, greater than 90 percent removal of As(V) was always
achieved. At lower coagulant dosages there was considerable scatter in the data attributed to poor
particle removal, high initial As(V) concentrations, and possible interferences from other anions in
the different waters tested.
2-3
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Effect of Coagulation pH
Sorg and Logsdon (1978) demonstrated that arsenic removal with alum coagulation is most
effective at pH 5 to 7, and ferric coagulation is most effective at pH 5 to 8. As discussed earlier,
Edwards (1994) summarized that at significant coagulant dosages As(V) removal was similar for
both alum and ferric coagulants at pH 7.6 or lower. At pH values greater than 7.6, however, the
average removals were 87 percent for 10 mg/L ferric chloride and only 61 percent for 20 mg/L alum.
Analyzing previously collected research data for As(III) removal by iron and aluminum
coagulation, Edwards (1994) demonstrated that removal of As (III) is much higher during iron
coagulation when compared with that of alum. Furthermore, As(III) removal by adsorption onto
aluminum hydroxides decreases markedly above pH 8.0.
Hering et al. (1996) observed the opposite effect. In coagulation experiments with ferric
chloride over the pH range of 4 to 9, pH did not appear to influence the As(V) removal. However,
strong pH dependence was observed for As(III) in coagulation experiments with ferric chloride, with
a minimum in removal efficiency at pH 6.0.
Logsdon et al. (1974) conducted several jar tests on spiked well water to analyze the initial
concentration and form of arsenic, and determine the type of coagulant most effective in arsenic
removal. The study found the initial arsenic concentration to have a significant effect on
removals. For initial As(V) concentrations between 0.1 and 1.0 mg/L, a dosage of 30 mg/L of
either alum or ferric sulfate hi the optimum pH range removed over 95 percent As(V). Above
an initial concentration of 1.0 mg/L, removals decrease with increasing concentrations. For
concentrations of As(in) greater than 0.1 mg/L, neither alum nor ferric sulfate dosed at 30 mg/L
could remove As(m) to concentrations below 0.05 mg/L. In both cas.es, higher coagulant dosages
(60 to 100 mg/L) resulted in higher removals.
Hering et al. (1996) demonstrated in coagulation experiments;, with ferric chloride dose of
4.9 mg/L at pH 7.0 and varied initial arsenic concentration from 2 to 100 ug/L, that both As(III) and
As(V) removal was independent of initial concentration. Cheng et al. (1994) showed that As(V)
2-4
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removal was independent of initial concentration when treated with 20 mg/L of alum and 30 mg/L
of ferric chloride while varying the initial As(V) concentration from 2.2 to 128
Effect of Co-occurring Inorganic Solutes
Co-occurring inorganic solutes, such as sulfate and calcium, may compete for surface binding
sites onto oxide surfaces and influence the adsorption of trace contaminants, such as arsenic. Hering
et al. (1996) investigated the effects of sulfate and calcium on the efficiency of As(III) and As(V)
removal during coagulation with 4.9 mg/L of ferric chloride. The results indicated that at pH below
7.0, As(III) removal was significantly decreased in the presence of sulfate. However, only a slight
decrease in As(V) was observed. At higher pH, removal of As(V) was increased in the presence of
calcium.
McNeill and Edwards (1997a) developed a simple model for predicting As(V) concentration
during coagulation with alum or ferric salts. Using inputs of aluminum hydroxide formed, ferric
hydroxide present in the influent, ferric hydroxide formed, and a single sorption constant, the model
predicted As(V) removal to within 13% for the 25 utility sampling events in this study. The authors
suggested an optimization hierarchy strategy for coagulation/filtration facilities which are unable to
meet arsenic removal requirements with their existing treatment scheme. If any As(III) is present
in the raw water, the most cost-effective method of improving removal is to convert poorly sorbed
As(III) to A;3(V). Thereafter, for facilities practicing alum coagulation, it is critical to minimize
residual soluble aluminum to enhance the formation of aluminum hydroxide solids which mediate
the As(V) removal. Jar testing should be performed to identify pH and coagulant dosage that might
be altered to reduce aluminum residuals. The final option is to increase the coagulant dosage or to
consider changing the coagulant type.
Summary
Coagulation is a successful technology for achieving As(V) removals greater than 90 percent.
Arsenic in the pentavalant arsenate form is more readily removed than the trivalent arsenite form.
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At pH 7.6 or lower iron and aluminum coagulants are of equal effectiveness in removing As(V).
However, iron coagulants are advantageous if pH is above 7.6, if soluble coagulant metal residuals
are problematic, or if As(III) is present in the raw water. In general, higher arsenic removal
efficiencies are achieved with increased coagulant dosages. The effectiveness of iron coagulants in
removing As(III) diminishes at pH 6.0. Recent studies have shown that arsenic removal is
independent of initial concentration. This contradicts initial findings which indicate that arsenic
removals decrease with increasing initial concentrations. Presence of sulfate significantly decreases
As(III) removal, whereas sulfate slightly affects As(V) removal. At pH higher than 7.0, removal of
As(V) increases in the presence of calcium.
2.2.2 Iron/Manganese Oxidation
Iron/Manganese (Fe/Mn) oxidation is dominant in facilities treating groundwater. Oxidation
to remove iron and manganese leads to formation of hydroxides that remove soluble arsenic by
precipitation or adsorption reactions.
Arsenic removal during iron precipitation is expected to be fairly efficient (Edwards, 1994).
Removal of 2 mg/L of iron achieved a 92.5 percent removal of As(V) from a 10 jj,g/L As(V) initial
concentration by adsorption alone. Even removal of 1 mg/L of iron is capable of adsorbing 83
percent of a 22 ug/L As(V) influent concentration. However, removal, of arsenic during manganese
precipitation is relatively ineffective when compared to iron even when removal by both adsorption
and coprecipitation are considered. For instance, precipitation of 3 mg/L manganese removed only
69 percent of As(V) of a 12.5 ug/L As(V) influent concentration.
Effect of Co-occurring Inorganic Solutes
McNeill and Edwards (1995) demonstrated that a Fe/Mn facility with 400 mg/L sulfate and
5.2 ug/L arsenic in the raw water attained 83 percent removal of arsenic. Results from two other
Fe/Mn facilities with 10 mg/L sulfate in the raw water showed 87 and. 93 percent arsenic removals.
This analysis suggests that sulfate interferes only slightly with sorption of arsenic onto ferric iron
precipitates.
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2.2.3 Coagulation Assisted Microfiltration
Arsenic is removed effectively during the coagulation process, as described in section 2.2.1.
Microfiltration is used as a membrane separation process to remove particulates, turbidity, and
microorganisms. In coagulation assisted microfiltration technology, microfiltration is used similarly
to a conventional gravity filter. The advantages of microfiltration over conventional filtration are
outlined below (Muilenberg, 1997):
• more effective microorganism barrier during coagulation process upsets;
• smaller floe sizes can be removed (smaller amounts of coagulants are required); and
• increased total plant capacity.
Vickers et al. (1997) reported that microfiltration exhibited excellent arsenic removal
capability. Addition of a coagulant did not significantly affect the membrane cleaning interval,
although the solids level to the membrane system increased substantially. With an iron and
manganese removal system, it is critical that all of the iron and manganese be fully oxidized before
they reach the membrane to prevent fouling (Muilenberg, 1997).
2.2.4 Enhanced Coagulation
The Disinfectant/Disinfection Byproduct (D/DBP) Rule requires the use of enhanced
coagulation treatment technique for the reduction of disinfection byproduct (DBF) precursors for
surface water systems which have sedimentation capabilities. This treatment technique involves
modifications to the existing coagulation process such as increasing the coagulant dosage, reducing
the pH, or both.
Cheng et al. (1994) conducted bench, pilot, and demonstration scale studies to examine
As(V) removals during enhanced coagulation. The enhanced coagulation conditions in these studies
included increase of alum and ferric chloride coagulant dosage from 10 to 30 mg/L, decrease of pH
from 7 to 5.5, or both. Results from these studies indicated the following:
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• Greater than 90 percent As(V) removal can be achieved under enhanced coagulation
conditions. As(V) removals greater than 90 percent were easily attained under all conditions
when ferric chloride was used.
• Enhanced coagulation using ferric salts is more effective for arsenic removal than enhanced
coagulation using alum. With an influent arsenic concentration of 5 ng/L, ferric chloride
achieved 96 percent As(V) removal with a dosage of 10 mg/L and no acid addition. When
alum was used, 90 percent As(V) removal could not be achieved without reducing the pH.
• Lowering pH during enhanced coagulation improved arsenic removal by alum coagulation.
With ferric coagulation pH does not have a significant effect between 5.5 and 7.0.
2.2.5 Lime Softening
Hardness is predominantly caused by calcium and magnesium compounds in solution.
Lime softening (LS) removes this hardness by creating a shift in the carbonate equilibrium. The
addition of lime to water raises the pH. Bicarbonate is converted to carbonate as the pH
increases, and as a result, calcium is precipitated as calcium carbonate. Soda ash (sodium
carbonate) is added if insufficient bicarbonate is present in the water to remove hardness to the
desired level. Softening for calcium removal is typically accomplished at a pH range of 9 to 9.5.
For magnesium removal, excess lime is added beyond the point of calcium carbonate precipita-
tion. Magnesium hydroxide precipitates at pH levels greater than 10.5. Neutralization is required
if the pH of the softened water is excessively high (above 9.5) for potable use. The most common
form of pH adjustment in softening plants is recarbonation with carbon dioxide.
LS has been widely used hi the U.S. for reducing hardness hi large water treatment
systems. LS, excess lime treatment, split lime treatment, and lime-soda softening are all common
in municipal water systems. All of these: treatment methods are effective hi reducing arsenic.
As(IH) or As(V) removal by LS is pH dependent. Oxidation of As(in) to As(V) prior to LS
treatment will increase removal efficiencies if As(ni) is the predominant form. Considerable
amounts of sludge are produced in a LS system and its disposal is expensive. Large capacity
systems may find it economically feasible to install recalculation equipment to recover and reuse
the lime sludge and reduce disposal problems. Construction of a new LS plant for the removal
of arsenic would not generally be recommended unless hardness must also be reduced.
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McNeill and Edwards (1997b) showed that the percentage of As(V) removal by calcium
carbonate and magnesium hydroxide is constant regardless of the initial As(V) concentration. At
pH 10.5-12, As(V) removal was 23 ±4 percent for removal by calcium carbonate over the range
of As(V) concentrations of 5-75 ^g/L. At pH 11, As(V) removal was 37 ±5 percent for removal
by magnesium hydroxide over the range of As(V) concentrations of 5-160 ^tg/L.
These results differ from those of Logsdon et al. (1974) who found that arsenic removal
was dependent on the initial arsenic concentration. In the optimum pH range, As(V) or oxidized
As(III) was reduced to 0.05 mg/L when the initial concentration was 0.35 mg/L or lower, while
As(in) was reduced to 0.05 mg/L when the initial concentration was less than 0.1 mg/L.
McNeill and Edwards (1997b) also found that As(V) removal by manganese hydroxide
solids is sensitive to As(V) initial concentrations. At pH of 10.5, there was about 80 percent
removal in the system with 75 jtg/L of As(V) versus about 30 percent of removal in the 150 /*g/L
As(V) solution.
Effect of Arsenic Oxidation State
As(V) was generally more effectively removed by LS than As(III). Sorg and Logsdon
(1978) conducted several LS pilot studies for the removal of both As(ni) and As(V). Two of the
tests were performed at pH 9.5 and 11.3. At a pH of 11.3, 99 percent of an initial As(V)
concentration of 0.58 mg/L was removed, whereas only 71 percent of an initial As(m) concentra-
tion of 0.34 mg/L was removed. At a pH of 9.5, 53 percent of an initial As(V) concentration of
0.42 mg/L was removed, whereas only 24 percent of an initial As(in) concentration of 0.24 mg/L
was removed.
EffecLoLpH
The optimum pH for As(V) removal by LS is approximately 10.5, and the optimum pH
for As(III) removal is approximately 11 (Logsdon, et al., 1974; Sorg and Logsdon, 1978).
Logsdon, et al. (1974) studied the effectiveness of excess LS on the removal of arsenic in jar tests.
The test water was a well water that contained 300 mg/L hardness as CaCO3 spiked with 0.4
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mg/L As(V). The pH varied between 8.5 and 11.5. At pH 10.5 and above, nearly 100 percent
arsenic removal was obtained. Below the optimum pH, the removals decreased with decreasing
pH. When the water was spiked with As (HI), removals were only around 75 percent in the
optimum pH range. Below the optimum pH range, removals sharply decreased to less than 20
percent. Removals of oxidized As(m), however, were almost identical to removals of As(V).
Arsenate removal during softening is controlled by formation of three solids including
calcium carbonate, magnesium hydroxide, and ferric hydroxide. Calcium carbonate and magnesium
hydroxide are produced from reactions which remove hardness from water after addition of lime,
caustic soda, and soda ash. Ferric hydroxide can be formed by precipitation of iron naturally present
in treatment plant influent or by addition of iron coagulant during sofi :ening.
A survey of full-scale plants by McNeill and Edwards (1995) indicated that soluble As(V)
removal is mediated primarily by sorption to magnesium and/or ferric hydroxide solids during water
softening operations. At softening facilities precipitating only calcite, soluble As(V) removal was
between 0 and 10 percent, whereas soluble As(V) removal at plants precipitating calcite and
magnesium and/or ferric hydroxide was between 60 and 95 percent.
McNeill and Edwards (1997b) performed bench-scale studies to investigate the role of iron
addition in optimizing the As(V) removal. At pH 9 without any iron addition, only a small amount
of As(V) was removed. However, adding increasing amounts of iron at this pH improved As(V)
removal, with 82 percent of the As(V) removed at an iron dose of 9 mg/L. At pH 9.7, a 38 percent
As(V) removal without iron addition was observed, versus 63±8.4 percent removal for iron dosages
between 0.25 and 9 mg/L.
Effect of Other Constituents
The competitive effects of sulfate and carbonate for surface binding sites onto magnesium
hydroxide surfaces and the influence on the adsorption of arsenic was examined by McNeill and
Edwards (1998). These effects were investigated in experiments with preformed magnesium
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hydroxide by adding 20 mg/L Mg+2 and raising the pH to 12 after spiking the source water with 20
mg/L of As(V). Samples were collected as pH was incrementally lowered at ten minute intervals.
At pH 1 1 and above, no appreciable sulfate or carbonate interference was observed compared
to the control case. However, at pH 10 to 10.5, the system with carbonate exhibited significantly
lower As(V) removal (78 percent versus 96 percent hi the control and sulfate systems), and nearly
twice as much of the magnesium was measured as soluble (6.3 versus 3.3 mg/L). These results
suggest that carbonate is somehow increasing the concentration of Mg+2, leaving less solid available
for As(V) sorption.
McNeill and Edwards (1997b) investigated the interference of orthophosphate on As(V)
removal by softening. Softening of raw water containing 15 jag/L As(V) at pH 12 indicated greater
than 95 percent As(V) removal. After spiking raw water with 32 ug/L orthophosphate, As(V)
removal was slightly lower at intermediate pH values. Because the amount of calcium and
magnesium removed during softening with and without orthophosphate was nearly equal, it seems
that orthophosphate interferes with arsenic removal by competing for sorption sites.
McNeill and Edwards (1997b) developed a simple model for predicting As(V) during
softening. Using inputs of calcium carbonate, magnesium, and ferric hydroxide solid concentrations
formed during softening, the model can predict percentage As(V) removal.
McNeill and Edwards (1997b) suggested an optimization hierarchy strategy for softening
facilities which are unable to meet arsenic removal requirements with their existing treatment
scheme similar to optimization of coagulation hierarchy. If As(III) is present, the most cost-effective
method of improving arsenic removal is preoxidation of As(III) to As(V), since As(V) is more
readily removed by precipitation of calcium carbonate and magnesium and ferric hydroxide. For
facilities that are currently precipitating only calcium carbonate, addition of iron can dramatically
improve arsenic removal. A final option is to raise the softening pH in order to precipitate
magnesium hydroxide which strongly sorbs As(V). These removal trends should be quantitatively
confirmed with jar testing for optimizing arsenic removal.
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Summary
Softening is a successful technology for achieving greater than 90 percent As(V) removals.
Arsenic in the pentavalent arsenate form is more readily removed than the trivalent arsenite form.
The optimum pH for As(V) removal by softening is approximately 10.5 and the optimum pH of
As(III) is approximately 11.0. Recent studies have shown that As(V) removal is independent of
initial concentration. This contradicts initial findings which indicate that As(V) removal is a
function of initial concentration. As(III) removal appears to depend on initial concentration.
Facilities precipitating only calcium carbonate observed lower As(V) removals when compared to
facilities precipitating calcium carbonate and magnesium and ferric hydroxide. Addition of iron
improves As(V) removal. Presence of sulfate and carbonate in the raw water does not interfere with
As(V) removal at pH 11. As(V) removal, however, is reduced in the presence of carbonate at pH
10 to 10.5 and the presence of orthophosphate at pH less than 12.0.
2.3 ABSORPTIVE PROCESSES
2.3.1 Activated Alumina
Activated Alumina (AA) is a physical/chemical process by which ions in the feed water
are sorbed to the oxidized AA surface. AA is considered an adsorption process, although the
chemical reactions involved are actually an exchange of ions (AWWA, 1990). Activated alumina
is prepared through dehydration of A1(OH)3 at high temperatures, and consists of amorphous and
gamma alumina oxide (Clifford and Lin, 1995). AA is used in packed beds to remove
contaminants such as fluoride, arsenic, selenium, silica, and NOM. Feed water is continuously
passed through the bed to remove contaminants. The contaminant ions are exchanged with the
surface hydroxides on the alumina. When adsorption sites on the AA surface become filled, the
bed must be regenerated. Regeneration is accomplished through a sequence of rinsing with
regenerant, flushing with water, and neutralizing with acid. The regenerant is a strong base,
typically sodium hydroxide; the neutralizer is a strong acid, typically sulfuric acid.
Many studies have shown that AA is an effective treatment technique for arsenic removal.
Factors such as pH, arsenic oxidation state, competing ions, empty bed contact time (EBCT), and
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regeneration have significant effects on the removals achieved with AA. Other factors include
spent regenerant disposal, alumina disposal, and secondary water quality.
EffecLoLpH
pH may have significant effects on arsenic removal with AA. "A pH of 8.2 is significant
because it is the "zero point charge" for AA. Below this pH, AA has a net positive charge resulting
in a preference for adsorption of anions, including arsenic (AWWA, 1990). Acidic pH levels are
generally considered optimum for arsenic removal with AA, however, some studies have presented.
conflicting effects of pH.
Several researchers have shown optimum pH for arsenic removal to be in the range of 5.5
to 6.0 for tests conducted on synthetic waters (Singer and Clifford, 1981; Rosenblum and Clifford,
1984). Others have also found improved performance at lower pH levels. Simms and Azizian
(1997) found that incrementally lowering the pH from 7.5 to 6.0 increased the number of bed
volumes which could be treated by 2 to 12 times. Hathaway and Rubel (1987) reported that the
performance of AA for As(V) removal deteriorates as the pH increases from 6.0 to 9.0. Operating
at an As(V) removal of 50 percent and at a pH of 5.5, a column treated 15,500 bed volumes (BVs).
For the same level of As(V) removal, a column operating at pH 6.0 treated 13,391 BVs and a column
operating at a pH of 9.0 treated only 800 BVs. Column studies conducted by Clifford and Lin
(1985) also showed this trend. For a target arsenic effluent concentration of 0.05 mg/L, a column
operating at apH of 6.0 treated 8,760 BVs of water, but at pH of 7.3 the column treated only 1,944
BVs. In contrast to these results, Benjamin et al. (1998) found almost no dependence on pH level.
The authors conducted isotherm and column studies with AA to investigate the removals of As(V)
at pH 5.5, 7.0, and 8.5. Results indicated increasing pH from 5.5 to 8.5 had almost no effect on
sorption of As(V) on AA.
Effect of Arsenic Oxidation-State
Like nearly all other treatment technologies, the oxidation state of arsenic plays a large role
in its removal; As(V) is much more easily adsorbed than As(III). Frank et al. (1986) conducted two
column runs at.pH 6. The influent in one run was 0.1 mg/L As(V), and in the other 0.1 mg/L As(III).
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The column treating water containing As( V) treated about 23,400 BVs before the effluent levels
reached 0.05 mg/L. The other column showed a breakthrough of As(III) almost immediately and
treated only 300 BVs before 0.05 mg/L was reached in the effluent. Benjamin et al. (1998) found
that adsorption of As(V) was much faster than adsorption of As(III). The authors also showed that
sorption onto AA was relatively rapid during the first few hours of exposure and slower thereafter.
The ratio of As(V) adsorption densities at 2 and 24 hours was approximately 88 percent, whereas
the ratio of As(III) adsorption densities was approximately 60 percent.
EffectJiLCDjoipfilingJons
Like ion exchange processes, AA exhibits preference for some ions. Interestingly, AA tends
to have increased preference for ions which ion exchange does not. AA, however tends to be
specific for arsenic and is not as greatly affected by competing ions (A^WWA, 1990). As is indicated
by the general selectivity sequence shown below (Clifford and Lin, 1995), AA preferentially
adsorbs H2AsO4- [As(V)J overH3AsO3[As(III)]:
OH' > H2AsO4- > Si(OH)3O-> F> HSeO/ > TOC > SO42- > H3AsO3
Several studies have illustrated the effects of this selectivity, particularly those associated
with sulfate and chloride. Benjamin et al. (1998) found little effect produced by either sulfate or
chloride. Increasing sulfate from 0 to 100 mg/L had only a small impact on the sorption of As(V).
The presence of chloride also did not affect As(V) removal. The addition of organics, however, had
a much greater effect. The addition of 4 mg/L DOC reduced As(V) sorption onto AA by about 50
percent.
Clifford and Lin (1986) found significant effects of sulfate and total dissolved solids (TDS)
on adsorption. They found the addition of 360 mg/L of sulfate and almost 1,000 mg/L TDS
decreased the sorption of As(V) onto AA by approximately 50 percent compared to sorption from
deionized water. Rosenblum et al. (1984) also reported that sulfate and chloride significantly
reduced arsenic removal in AA systems. Arsenic removal in a water containing approximately 530
mg/L of chloride was 16 percent less than that achieved in a deioni/:ed water, and the presence of
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720 mg/L of sulfate resulted in more than 50 percent less arsenic removal than that achieved in
deionized water.
Simms and Azizian (1997) reported competition with silicate. In this particular study, the
AA media became saturated with silicate much more quickly than with arsenic. No de-sorption of
silicate was observed after saturation.
The operation of AA beds, and in particular the EBCT, can also play a role in arsenic
removal. EBCT represents the lenght of time in which the feed water is in contact with the AA
medium. Benjamin et al. (1998) conducted AA column tests using arsenic-spiked water from Lake
Washington. All the column tests were run by adjusting the feed solution to pH 7. Sampling ports
at various points in the system allowed EBCTs ranging from 2.5 to 15 minutes to be tested. Low
arsenic concentrations (i.e. <5 jig/L) were achieved for more than 2,000 hours of operation.
Comparing EBCTs, the data show that adsorption increased slightly with increasing EBCT.
Regeneration
Regeneration of AA beds is usually accomplished using a strong base solution, typically
concentrated NaOH. Relatively few BVs of regenerant are needed. After regeneration with strong
base, the AA medium must be neutralized using strong acid; typically two percent sulfuric acid.
Arsenic is more difficult to remove during regeneration than other ions such as fluoride (Clifford and
Lin, 1995). Because of this, slightly higher base concentrations are used; typically 4 percent NaOH.
Even at this increased concentration , however, not all arsenic may be eluted. Clifford and Lin
(1986) found only 50 to 70 percent of arsenic was removed from the AA columns during
regeneration. Other researchers have also documented the difficult regeneration of AA for arsenic.
Regeneration tests conducted by Benjamin et al. (1 998) indicated that exposure of the AA medium
to 0.1 N NaCl or 0.2 N NaOH did not regenerate the AA to a significant extent. Arsenic recovery
was limited and in most cases was less than 50 percent of the sorbed arsenic. Higher recoveries have
been reported, however. Hathaway and Rubel (1987) found that 80 percent of the adsorbed arsenic
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was eluted using 1.0 to 1.25 M NaOH solution. Simms and Azizian (1997) found that up to 85%
of the capacity of an AA bed could be recovered using NaOH.
Regeneration also affects successive bed life and efficiency. Bed life is shortened and
adsorption efficiency is decreased by regeneration. Benjamin et al. (1998) found that arsenic
breakthrough patterns from the AA columns using regenerated media were qualitatively similar to
those using fresh media, but the removal efficiency declined slightly after each of two regenerations.
Clifford (1986) demonstrated that regeneration has a clearly negative effect on the
adsorption capacity of activated alumina. The unrecovered As(V) and changes in the AA surface
induced by the regeneration process may cause the length of the adsoiption runs to decrease by 10
to 15 percent after each regeneration.
Resin Fouling
Much like ion exchange resin, AA media may be fouled. Fouling reduces the number of
adsorption sites thus decreasing removal effectiveness.
Hydraulic considerations should also be given. During treatment, AA media may become
clogged with suspended solids present in the feed water. This can result in increased headless across
the bed. If the headless buildup is significant, the media must be backwashed to removed the solids.
Simms and Azizian (1997) found that headloss buildup across the bed after 75,000 BVs treated was
minimal for a groundwater with 2 mg/L suspended solids and which was not pre-filtered.
In addition to suspended solids, Clifford and Lin (1995) note that silica and mica are
particularly problematic foulants. In a study performed in Hanford, California, mica fouling was
found to be a significant problem (Clifford and Lin, 1986).
Operational Considerations
Experience with AA processes is limited and full-scale applications are virtually non-
existent. Therefore a large amount of information still needs to be obtained. The operational
experience which has been developed, however, provides important information to be considered
for AA process; these are discussed here.
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AA beds may be operated in series or parallel. Series operation increases removal and helps
prevent leakage, but limits throughput (leakage simply refers to elevated levels of arsenic in the
effluent). Parallel operation on the other hand increases throughput, but does not improve effluent
quality (AWWA, 1990). When operated in series, a "merry-go-round" configuration is often used.
This configuration uses three beds: two in production and one in regeneration mode at a given time.
When exchange capacity of the first bed in series is exhausted, the first bed is removed from service
to be regenerated. The second bed in series then becomes the first and a fresh regenerated bed is
brought on-line to become the second. This allows the maximum exchange capacity of beds to be
used and prevents leakage since a fresh bed is always last in line. This also helps minimize
regeneration frequency.
Degradation of AA media must also be considered. Alumina tends to dissolve over
successive cycles due to the strong base/strong acid cycling during regeneration. As a result of this,
alumina beds may become "cemented" if close care is not given (EPA, 1994). Backwashing the AA
media may help prevent cementation. Another important consideration is operator involvement.
Strong acid and strong base are handled on a frequent basis and can present a safety hazard. An
operator must be capable of handling these chemicals and must have a good understanding of pre-
treatment, post-treatment, and regeneration practices if the process is to be operated efficiently. This
presents a problem particularly for small systems.
DJspxtsaLJssu.es
Disposal of both spent regenerant and spent media is an important issue with arsenic removal
using AA. Spent regenerant can contain high levels of arsenic. Simms and Azizian (1997)
documented 20 to 40 mg/L of arsenic in spent regenerant liquid. Although little work has been done
in this area, it has been speculated that the spent AA media would pass toxicity tests and could be
landfilled. It is doubtful if spent regenerant could be discharge directly to the a sanitary sewer.
There is evidence, though, that spent regenerant may be effectively treated prior to disposal
(AWWA, 1990). This is possible because, during regeneration and acidification, enough aluminum
dissolves to make precipitation of A1(OH)3 a potential treatment. Arsenic is removed through its
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coprecipitation with the solid aluminum hydroxide. Arsenic is removed via the aluminum sludge
which can be subsequently dried and landfilled if toxicity limits are not exceeded.
Although the possibility of regenerarit reuse exists, it may not be feasible for arsenic removal.
Direct reuse would probably not be possible due to the strong affinity of AA for arsenic. In other
words, arsenic in the reused regenerant may actually be added to the column during regeneration.
Spent regenerant, however, may be treated prior to reuse. By precipitating the arsenic from the
regenerant, reuse may be possible assuming the regenerant solution was replenished and remained
concentrated enough to replenish the AA bed.
Secon da.ry Effects
AA processes will produce changes to the effluent water quality (EPA, 1994). Because pre-
treatment is typically used to reduce the pH to low levels (less than 6.0) to optimize the process, the
effluent pH will be less than typically desired in the distribution system. For this reason, post-
treatment to raise the pH would be necessary. Another important effect of AA is the increased TDS
levels of the effluent. AA processes tend to increase TDS levels which may have important
implications for some utilities, such as corrosion issues or water quality issues.
2.3.2 Iron Oxide Coated Sand
Iron oxide coated sand (IOCS) is a tare process which has sho\m some tendency for arsenic
removal. IOCS consists of sand grains coated with ferric hydroxide which are used in fixed bed
reactors to remove various dissolved metal species. The metal ions are exchanged with the surface
hydroxides on the IOCS. IOCS exhibits selectivity in the adsorption and exchange of ions present
in the water. Like other processes, when the bed is exhausted it must be regenerated by a sequence
of operations consisting of rinsing with tegenerant, flushing with water, and neutralizing with
strong acid. Sodium hydroxide is the most common regenerant and sulfuric acid the most
common neutralizer.
Several studies have shown that IOCS is effective for arsenic removal. Factors such as
pH, arsenic oxidation state, competing ions, EBCT, and regeneration have significant effects on
the removals achieved with IOCS.
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Effect of pH
pH appears to have an effect on arsenic adsorption by IOCS. Benjamin et al. (1998)
conducted isotherm and column studies with IOCS to investigate the removals of As(V) at various
pH levels. Results indicated that increasing the pH from 5.5 to 8.5 decreased the sorption of As(V)
by approximately 30 percent.
EHeci_olArseni^OxidatiojiJStaJe
As with other processes, the oxidation state of arsenic plays a role in its removal: As(V)
appears to be more easily removed than As(III). Benjamin et al. (1998) showed that As(V) sorption
onto IOCS was much more rapid than As(III) sorption during the first few hours of exposure and
slower thereafter. The ratio of As(V) adsorption densities at 2 and 24 hours was approximately 60
percent, whereas the ratio of As(III) adsorption densities was only about 50 percent.
Concentrations of competing ions will be an important consideration for arsenic removal with
IOCS. Benjamin et al. (1998) evaluated the effect of sulfate and chloride on IOCS arsenic
adsorption. They found that increasing sulfate from 0 to 100 mg/L had only slight impact on the
sorption of As(V), and the presence of chloride did not appear to affect As(V) removal. Organic
matter, however, did appear to present some competition for arsenic. The addition of 4 mg/L DOC
reduced As(V) sorption by about 50 percent.
EJBfecliifJEinpJtyJBed-C^
The EBCT can affect the arsenic removal efficiency of IOCS. Benjamin et al. (1998)
conducted continuous flow IOCS column tests using arsenic-spiked water from Lake Washington.
All tests were run by adjusting the feed solution to pH 7. Sampling ports at various points in the
system allowed EBCTs ranging from 2.5 to 15 minutes to be tested. Low arsenic concentrations (i.e.
<5 ng/L) were achieved for more than 2,000 hours of operation. Adsorption seemed to increase
slightly with increasing EBCT. Based on the adsorption density at complete breakthrough, the initial
capacity of the IOCS for either As(V) or As(III) was between 175 and 200 jig As/mL of media.
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Regeneration
Regeneration of IOCS is performed in a similar fashion to that performed with activated
alumina processes. Regeneration is accomplished using a strong; base, typically NaOH, and
subsequent neutralization is accomplished using strong acid, typically H2SO4. Regeneration tests
conducted by Benjamin et al. (1998) indicated that exposure of the IOCS medium to 0.1 N NaCl or
0.2 N NaOH did not regenerate IOCS to a significant extent. Arsenic recovery was limited and in
most cases was less than 50 percent of the sorbed arsenic. The arsenic breakthrough patterns from
the IOCS columns using regenerated media were qualitatively similar to those using fresh media,
but the removal efficiency declined slightly after each of two regeneration steps.
2.4 ION EXCHANGE
2.4.1 Introduction
Ion exchange (IX) is a physical/chemical process by which an ion on the solid phase is
exchanged for an ion in the feed water. This solid phase is typically a synthetic resin which has been
chosen to preferentially adsorb the particular contaminant of concern. To accomplish mis exchange
of ions, feed water is continuously passed 'through a bed of ion exchange resin beads in a downflow
or upflow mode until the resin is exhausted. Exhaustion occurs when all sites on the resin beads
have been filled by contaminant ions. At this point, the bed is regenerated by rinsing the DC column
with a regenerant - a concentrated solution of ions initially exchanged from the resin. The number
of bed volumes that can be treated before exhaustion varies with resin type and influent water
quality. Typically from 300 to 60,000 BVs can be treated before regeneration is required. In most
cases, regeneration of the bed can be accomplished with only 1 to 5 EtVs of regenerant followed by
2 to 20 BVs of rinse water.
Important considerations in the applicability of the IX process; for removal of a contaminant
include water quality parameters such as pH, competing ions, resin type, alkalinity, and influent
arsenic concentration. Other factors include the affinity of the resin for the contaminant, spent reg-
enerant and resin disposal requirements, secondary water quality effects, and design operating
parameters.
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2.4.2 Effect of pH
The chloride-arsenate exchange chemical reaction typically occurs in the range of pH 8 to
9 when using chloride-form, strong-base resins (Clifford and Lin, 1995). IX removals with strong-
base resins, though, is typically not sensitive to pH in the range of pH 6.5 to 9.0 (Clifford, et al.,
1998). Outside of this range, however, arsenic removal decreases quickly. Groundwaters which are
naturally contaminated with arsenic typically exhibit fairly high pH, giving DC a slight advantage
for these types of source water. Adjustment of pH prior to IX for arsenic removal is generally not
necessary.
2.4.3 Effect of Competing Ions
Competition from background ions for IX sites can greatly affect the efficiency, as well as
the economics, of IX systems. The level of these background contaminants may determine the
applicability of IX at a particular site. Typically, strong-base anion exchange resins are used hi
arsenic removal. Strong-base anion resins tend to be more effective over a larger range of pH than
weak-base resins. The order of exchange for most strong-base resins is given below, with the
adsorption preference being greatest for the constituents on the far left.
HCrCv > CrO42' > ClCv > SeO42-> SO42' > NCy > Br > (HPO42-, HAsO42-, SeO32-, CO32-) > CN'
> NO2- > Cl- > (HjPO4-, H2AsO4-, HCCy) > OH" > CH3COQ- > F'
These resins have a relatively high affinity for arsenic in the arsenate form (HAsO42~),
however, previous studies have shown that high TDS and sulfate levels compete with arsenate and
can reduce removal efficiency (AWWA, 1 990). In general, ion exchange for arsenic removal is only
applicable for low-TDS, low-sulfate source waters. Source waters with TDS levels above 500 mg/L
and sulfate levels above 25 mg/L are not recommended. Previous studies have confirmed this
generalization; the low-sulfate/low-TDS source water in a Hanford, CA study proved to be amenable
to DC treatment whereas the high-sulfate/high-TDS source water in a San Ysidro, NM study proved
to be impractical for IX treatment (Clifford and Lin, 1986; Clifford and Lin, 1995).
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If nitrate removal is being performed concurrent with arsenic removal, sulfate level can also
be an important factor in arsenic removal. Clifford and others (1998) have shown that when sulfate
levels are low (about 40 mg/L), the number of BVs to exhaustion is limited by nitrate breakthrough.
If the sulfate level is high (about 100 mg/L), however, the number of BVs to exhaustion is limited
by arsenic breakthrough. In other words, sulfate competes with both nitrate and arsenic, but
competes more aggressively with arsenic than nitrate.
The presence of iron, Fe(III), in feed water can also affect arsenic removal. When Fe(III) is
present, arsenic may form complexes with iron. These complexes are not removed by IX resins and
therefore arsenic is not removed. Utilities with source waters high in Fe(III) may need to address
this issue for IX use or evaluate other treatment techniques for arsenic removal (Clifford, et al.,
1998).
When an ion is preferred over arsenate, higher arsenic levels in the product water than exist
in the feed water can be produced. If a resin prefers sulfate over arsenate, for example, sulfate ions
may displace previously sorbed arsenate ions, resulting in levels of arsenic in the effluent which are
greater than the arsenic level in the influent. This is often referred to as chromatographic peaking.
As a result, the bed must be monitored and regenerated well in advance of the onset of this peaking.
Clifford and Lin (1995) recommend operating the bed to a known BV setpoint to avoid peaking.
2.4.4 Resin Type
As stated earlier, strong-base resins are typically used in IX arsenic removal. These resins,
however, tend to prefer some ions, sulfate and chloride in particular, over arsenate. As mentioned
above, this can result in chromatographic: peaking if beds are not monitored adequately. Recent
studies have also found that sulfate-selective resins tend to be superior to nitrate-selective resins for
arsenic removal (Clifford, et al., 1998). Future research, howevesr, may produce monovalent-
selective resins which will be arsenate-selective and may eliminate non-arsenic ion competition
(EPA, 1994).
Many resins are available for arsenic removal. Some of the commercially available resins
which have been used in relevant IX studies are summarized in Table 2-1. Data in Table 2-1
represent BVs to exhaustion using virgin IX resins. It should be noted, however, that the capacity
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of the bed may decrease slightly over time. Choice of resin will ultimately be site-specific, making
preliminary studies a necessity to determine optimum resin type.
2.4.5 Process Configuration
Properly configuring IX columns can improve arsenic removal and help minimize
regeneration frequency. This is because arsenic "leakage" often occurs in IX columns. In some
situations, series operation or implementation of certain operating methods may be needed to achieve
low arsenic levels.
Series operation, also known as "merry-go-round" operation, uses three beds: two in
production and one in regeneration mode at a given time. When exchange capacity of the first bed
in series is exhausted, the first bed is removed from service to be regenerated. The second bed in
series then becomes the first and a fresh regenerated bed is brought on-line to become the second.
This allows the maximum exchange capacity of beds to be used and prevents leakage since a fresh
bed is always last in line. This also helps minimize regeneration frequency (EPA, 1995).
Another approach for niinimizing effluent levels is to operate IX columns in "counter-current
flow" operation. In this mode, feed water is applied in one direction (e.g., downward) and the
regenerant is applied in the opposite direction (e.g., upward). This minimizes leakage from the
column. Typically columns are designed for "co-current flow" operation where the feed water and
regenerant are applied in the same direction. Co-current operation increases chances for leakage,
however, since regeneration in this mode concentrates the contaminant on the effluent end of the IX
column. Using the "counter-current flow" method also minimizes regenerant requirements, i.e.
volume and concentration (EPA, 1995).
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TABLE 2-1
Typical IX Resins for Arsenic Removal
Resin
Dowex 1 1
lonac ASB-2
Dowex SBR-1
lonac ASB-1
and
Dowex 1 1
• A-300E
(bench-scale)
• A-300E
(full-scale)
Type
2
2
1
1
2
Operating Parameters
Bed Volume: 0.8 cu ft
Flowrate: 1 gpm
EBCT: 5.6 min
Depth: 2.5 - 5 ft
Sulfate/As Ratio: 60:1
TDS/As Ratio: 2500:1
Bed Volume: 0.8 cu ft
Flowrate: 1 gpm
EBCT: 5.6 min
Depth: 2.5 - 5 ft
Sulfate/As Ratio: 60:1
TDS/As Ratio: 2500:1
Bed Volume: l.Ocuft
Flowrate: 1 gpm
EBCT: 7.5 min
Depth: 3.8 ft
Sulfate/As Ratio: NR
TDS/As Ratio: NR
Bed Volume: 0.014 cu ft
Flowrate: NR
EBCT: NR
Depth: NR
Sulfate/As Ratio: 420:1
TDS/As Ratio: 9200:1
Bed Volume: 0.001 8 cu ft
Flowrate: 0.035 gptn
EBCT: NR
Depth: 1.33 ft
Sulfate/As Ratio: 300:1
TDS/As Ratio: NR
BVs to Exhaustion
4,200
-
4,940
2,800
• 200
• 400-500 (projected
if oxidation to
As(V)is
preformed)
• 1,340-1,640
• 5,000-7,000'
Reference
Clifford and
Lin (1986)
Hathaway
and Rubel
(1987)
Fox (1989)
Clifford and
Lin
(1985)
Malcolm
Pirnie (1992)
NR = Not Reported
2.4.6 Secondary Effects
Chloride-form resins are often used in arsenic removal. Chloride ions are displaced from the
column as contaminants (arsenic) are sorbed onto the column. As a result, the potential exists for
increases in the chloride concentration of the product water. Increases in chlorides can greatly
increase the corrosivity of the product water. Chlorides increase the corrosion potential of iron and
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as a result increase the potential for red water problems (EPA, 1995). Corrosion problems are
worsened when high chloride levels are intermittent. In situations where chlorides pose a problem,
demineralization, blending, or alternate treatment techniques may be required.
Also, effluent pH may be lowered as a result of IX treatment. pH of the product water may
be less than 7 at the beginning of a cycle. Again, decreases in pH may increase the corrosivity of
the effluent. In some situations, pH restabilization may be necessary to prevent disturbances in the
distribution system.
2.4.7 Resin Fouling
IX resin beads may be fouled if appropriate pretreatment is not practiced. Generally, fouling
of IX resins is caused by scaling of minerals (i.e. Ca) or by particulates in the feed stream. Iron
precipitates have also been known to cause resin fouling (Malcolm Pirnie, 1993a). If scaling is a
problem, chemical addition may be needed to lower the scale-forming potential of the feed water.
If suspended solids are found in the feed stream, multi-media filtration ahead of IX columns may
be necessary. A previous study performed in Hanford, California found that IX resin was
significantly fouled by mica present in the source water. This was indicated by a 3-5 percent
decrease in total B Vs to exhaustion over consecutive cycles, and by a black coating on the exhausted
resin. Most, but not all, of the black coating could be removed from the resin beads during the NaCl
regeneration cycle (Clifford and Lin, 1986).
2.4.8 Regeneration
With chloride-form resins, concentrated NaCl solution is typically used as the regenerant.
Only a few number of BVs of regenerant are usually required to replenish the resin, depending on
the solution strength. Arsenic elutes readily from IX columns, regardless of resin type, mainly
because it is a divalent ion and as such is subject to selectivity reversal in high ionic strength (> 1M)
solution (Clifford and Lin, 1995). Clifford and Lin also found that dilute regenerants tend to be
more efficient than concentrated regenerants in terms of the ratio of regenerant equivalents to resin
equivalents. For example, they found that two resins (Dowex-11 and lonac ASB-2) could be
regenerated equivalently using either 2 BVs of 1.0 N NaCl or 5 BVs of 0.25 N NaCl in "co-current
2-25
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flow" operation. Also, a rinsing cycle is required after regeneration; typically only a few BVs are
required for rinsing as well.
2.4.9 Regenerant Reuse and Treatment
Spent regenerant is produced during IX bed regeneration. Typically this spent regenerant
will have high concentrations of arsenic and other sorbed contaminants. Spent regenerant must be
treated and/or disposed of appropriately. This can be an expensive part of the IX process and must
be given thorough consideration. Spent brine can be disposed of either directly to a surface water
source, or indirectly to a sanitary sewer, depending on contaminant levels. Spent solution can also
be treated and disposed of as described below. Spent regenerant, however, may be reused many
times. Clifford and others (1998) estimate; that regenerants may be used 25 times or more before
treatment and disposal are required. Regenerants do not need treatment prior to reuse, except to
replenish the chloride concentration to maintain a 1 M solution. Once the contaminant concentration
becomes too high in the regenerant, the spent solution must be treated and/or disposed.
Treatment of spent regenerant is accomplished in a number of ways. First, the spent
regenerant can be dewatered in some fashion. Common methods of dewatering IX residuals include
mechanical dewatering, drying beds, gravity thickeners, and lagoon dewatering. The solids
generated by these processes would need to be tested for toxiciry and disposed accordingly. If
determined to be non-toxic according to disposal regulations, the dried solids could be landfilled.
Waste liquid generated by these drying processes could be either directly discharged to a surface
water source or indirectly discharged to the sanitary sewer, depending on contaminant levels
(Malcolm Pimie, 1996). Second, spent brine can be treated by precipitation. Clifford and Lin
(1995) have shown that arsenic levels can be substantially redua^d using iron and aluminum
coagulants as well as lime. Much greater than the stoichiometric amounts (up to 20 times as much),
however, are needed in actual practice to reduce arsenic to low levels. In addition, pH adjustment
may be necessary to ensure optimum coagulation conditions. Reductions from 90 mg As(V)/L to
less than 1.5 mg As(V)/L have been seen using iron and aluminum metal salts (Clifford and Lin,
1995). Both coagulant types seem to work well, however, iron precipitates tend to settle better due
to their weight. Dried sludge from brine reduced to 1.5 mg As(V)/L using precipitation passed an
EP toxicity test with only 1.5 mg/L As(V) in the leachate. In this situation, dried sludge could have
2-26
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been disposed of in a landfill. Reuse of decontaminated regenerant has yet to be evaluated, however,
the possibility of reuse does exist.
2.4.10 EBCT
A few studies have been performed to test the effect of EBCT on IX performance.
Clifford and Lin (1986) reduced EBCT from 5 to 1.4 in a Hanford, CA study and found no
significant reduction hi arsenic removal performance. In a recent AWWARF study, four IX
columns were run with EBCTs varying between 2.5 and 15 minutes. Data from this study show
that the shorter the EBCT, the more BVs can be treated before breakthrough. The disadvantage
to shorter EBCT, however, is increased regeneration frequency. Based on these data, shorter
EBCTs may be preferred to reduce capital costs (AWWARF, 1998).
2.4.11 Typical Design Parameters
Through extensive research, Clifford and others (1998) assembled typical operating
parameters and suggested options for ion exchange processes. Although many design parameters
should be tailored to the specific treatment situation, Table 2-2 gives typical values and options.
TABLE 2-2
Typical Operating Parameters and Options for IX
1.5 minute EBCT (15 gpm/ft2 at 3 ft/day)
0.5 - 1.0 M NaCl (1-2 eq ClVeq resin)
Operate the column to a fixed BV endpoint (to prevent leakage)
Regenerant Surface Loading Velocity should be greater than 2 cm/min
Regenerant may be used 25 times or more (with Cl~ concentration of 1 M
maintained)
Ferric coagulant should be used for Fe(OH)3»As from regenerant waste
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2.5 MEMBRANE PROCESSES
2.5.1 Introduction
Membranes are a selective barrier, allowing some constituents to pass while blocking the
passage of others. The movement of constituents across a membrane requires a driving force (i.e.
a potential difference between the two sides of the membrane). Membrane processes are often
classified by the type of driving force, including pressure, concentration, electrical potential, and
temperature. The processes discussed here include only pressure-driven and electrical potential-
driven types.
Pressure-driven membrane processes are often classified by pore size into four categories:
microfiltration (MF), ultrafiltration (UF), nanofiltration (NF), and reverse osmosis (RO). Typical
pore size classification ranges are given in Figure 2-1. High-pressure processes (i.e., NF and RO)
have a relatively small pore size compared to low-pressure processes (i.e., MF and UF). Typical
pressure ranges for these processes are given in Table 2-3. NF and RO primarily remove
constituents
through chemical diffusion (Aptel and Buckley, 1996). MF and UF primarily remove constituents
through physical sieving. An advantage of high-pressure processes is that they tend to remove
a broader range of constituents than low-pressure processes. However, the drawback to broader
removal is the increase in energy required for high-pressure processes.
TABLE 2-3
Typical Pressure Ranges for Membrane Processes
Membrane Process
MF
UF
NF
RO
Pressure Range
5-45 psi
7 - 100 psi
50 - 150 psi
100 - 150 psi
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SEPARATION SPECTBWM
Mtcn>m«ter» (Log Seato)
Appro. MoMcular w.lgnt
Relative
Size of
Common
Materials
Desalination
Processes
Filtration
Processes
o.oc
1OO 2C
F s
f Aqueous S
f Metal Ion \
Atomic 1
Radius 1
t 00
0 1000 20,<
t
I Pyroge
r Aquatic OOC '
"9ar ^
alt ^
Li!
k cJ
1 0.
oo 100,000
Virus
' k
k *
>umln Protein J
oldal Slllc* j
REVERSE OSMOSIS 1
f NANOFILTRATION J
a^r1
| ULTRAFILTRATION
1
1
soo.ooo
J
k
*b«toc J
>
I 1
I Ciyptaa i
t sperMlunJ
1 Ola
j c
L Coal
Bacteria
) 10
rdla /
rst J
Dual i
1
a Pollen
D 10
_J
^ar — ^
•. MEDIA FILTRATION ~>
ar-^
f MICROF1LTRATIOM ^>
>o
Figure 2-1 Pressure Driven Membrane Process Classification
(Westerhoff and Chowdhury, 1996)
Electrical potential-driven membrane processes can also be used for arsenic removal.
These processes include, for the purposes of this document, only electrodialysis reversal (EDR).
In terms
of achievable contaminant removal, EDR is comparable to RO. The separation process used in
EDR, however, is ion exchange (Aptel and Buckley, 1996). EDR is discussed further in Section
2.5.8.
2.5.2 Important Factors for Membrane Performance
Commercial pressure-driven membranes are available in many types of material and in
various configurations. The chemistry of the membrane material, in particular surface charge and
hydrophobicity, play an important role in rejection characteristics since membranes can also remove
contaminants through adsorption. Membrane configuration and molecular weight cut-off (MWCO),
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i.e. pore size, also influence rejection properties, as well as operational properties, to a great extent.
These options must be chosen appropriately depending on source water characteristics and removal
requirements.
Source water quality is also important in the selection of a membrane process. Water quality
can have significant effects on membrane operation and rejection. Water temperature is very
important to all membrane processes. Lower water temperatures will decrease the flux at any given
pressure. To compensate, additional membrane area and/or higher feed pressures must be provided
to maintain equivalent production at lower temperatures. Depending on source water quality,
pretreatment is often necessary, particularly with the high-pressure processes. The small pore size
of NF and RO membranes makes them more prone to fouling than UF or MF membranes. The
application of NF and RO for surface water treatment is generally not accomplished without
extensive pretreatment for particle removal and possibly pretreatment for dissolved constituents.
The rejection of scale-causing ions, such as calcium, can lead to precipitation on the membrane
surface. Organic compounds and metal compounds, such as iron and manganese, can promote
fouling as well. Precipitation can result in irreversible fouling and must be avoided by appropriate
pretreatment, including addition of anti-scaling chemical and/or acid, to the feed water.
The percentage of product water mat can be produced from ttie feed water is known as the
recovery. Recovery for MF and UF is typically higher than recovery for RO and NF. The recovery
is limited by the characteristics of the feed water and membrane properties. Typical recoveries for
membrane processes are given in Table 2-4.
TABLE 2-4
Typical Recovery for Membrane Processes
Membrane Process
MF
UF
NF
RO
Recovery
to 99%
to 95%
to 85%
30-85%
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2.5.3 Arsenic Removal with Membrane Processes
Membrane processes can remove arsenic through filtration, electric repulsion, and adsorption
of arsenic-bearing compounds. If particulate arsenic compounds are larger than a given membrane
pore size, they will be rejected due to size exclusion. Size, however, is only one factor which
influences rejection. Studies have shown that some membranes can reject arsenic compounds which
are one to two orders of magnitude smaller than the membrane pore size, indicating removal
mechanisms other than just physical straining (AWWARF, 1998). Shape and chemical characteris-
tics of arsenic compounds play important roles in arsenic rejection. Membranes may also remove
arsenic compounds through repulsion by or adsorption on the membrane surface. These depend on
the chemical characteristics, particularly charge and hydrophobicity, of both the membrane material
and the feed water constituents. Many studies have been performed which evaluated various
membrane processes for arsenic removal. These processes and corresponding research are discussed
in the remainder of this section.
2.5.4 Microflltration
Microfiltration's viability as a technique for arsenic removal is highly dependent on the size
distribution of arsenic-bearing particles in the source water. MF pore size is too large to
substantially remove dissolved or colloidal arsenic. Although MF can remove particulate forms of
arsenic, this alone does not make the process efficient for arsenic removal unless a large percentage
of arsenic is found in this form. Arsenic found in groundwater is typically less than 10 percent
particulate while arsenic found in surface waters can vary from 0 percent to as much as 70 percent
particulate (AWWARF, 1998; McNeill and Edwards, 1997). Unfortunately, the percentage of
particulate arsenic does not seem to be related to specific water types. In a recent study, AWWARF
(1998) did not find arsenic size distribution to correlate with turbidity or organic content, indicating
that arsenic size distribution was specific to individual waters.
To increase removal efficiency in source waters with a low percentage of particulate arsenic
content, MF can be combined with coagulation processes. Coagulation assisted microfiltration for
arsenic removal is discussed in Section 2.2.3. For utilities using MF alone for particulate arsenic
removal, removal would primarily depend on the influent arsenic concentration and percentage of
particulate arsenic since the MF rejection mechanism is mechanical sieving. Therefore, the
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effectiveness of MF arsenic rejection is a function of pore size. Variation in MF performance is due
to pore size distribution.
2.5.5 Ultrafiltration
Ultrafiltration processes are generally capable of removing some colloidal and particulate
constituents, based upon the above discussion on particulate arsenic occurrence. Considering this,
UF alone, like MF, may not be a viable teclmique for arsenic removal for groundwaters, however,
UF may be appropriate for surface waters with high colloidal and particulate arsenic concentrations.
Recent research has found that electric repulsion of UF may play an important role in arsenic
rejection and increase rejection beyond that achievable with only p»ore size-dependent sieving.
AWWARF (1998) performed bench-scale tests on two low-MWCO UF membranes. Single element
testing was performed on Desal GM and FV UF membranes for a spiked, deionized water. Flat sheet
testing was also performed on Desal GM, FV, and PM UF membranes for spiked, deionized water.
Since the samples were spiked, no particulate or colloidal arsenic was present. Results of this study
are given in Table 2-5.
TABLE 2-5
As(V) and As(III) Removal by UF Membranes
Membrane Type
Single Element
GM2540F
GM2540F
GM2540F
GM2540F
FV2540F
FV2540F
Flat Sheet
GM
FV
PW
MWCO
8,000
8,000
8,000
8,000
10,000
10,000
8,000
10,000
10,000
Membrane
Charge
(-)
(-)
(-)
(-)
None
None
(-)
None
None
Arsenic
Species
V
V
in
m
V
ni
V
V
V
PH
6.9
2.0
7.2
10.8
6.9
6.8
Total Arsenic
Rejection (%)
63
8
<1
53
3
5
52
NA
5
NA: Not Available
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For the negatively charged GM2540F membrane, As(V) rejection was high at neutral pH
but very low at acidic pH. On the other hand, with the same membrane, As(m) rejection was
high at basic pH and negligible at neutral pH. The uncharged FV2540F membrane showed poor
rejection of both As(V) and As(in) at neutral pH. High rejection rates were seen even though the
MWCO of the membranes were two orders of magnitude larger than the arsenic compounds
(AWWARF, 1998). The authors theorize that the high rejection rates seen were due to
electrostatic interaction between the negatively charged membrane surface and the arsenic ions.
This will be pH dependent since the anionic As(V) and the nonionic As(ni) will be charged (-
protonated/deprotonated) at different pH levels. In effect, membrane charge and pH may play an
important role in arsenic rejection. In fact, the authors found that electrostatic repulsion becomes
increasingly important moving from RO to NF to UF, while size exclusion becomes increasingly
important moving from UF to NF to RO. The flat sheet testing produced rejection rates
comparable, and slightly conservative, to the single element rejection rate. As with single element
testing, the negatively charged membrane proved more effective for arsenic rejection than the
neutral charged membrane.
AWWARF (1998) also performed UF pilot-scale tests. Single element pilot tests were
performed on two groundwaters, one with a DOC level of 11 mg/L and one with a DOC level of
1 mg/L, and a spiked, finished surface water. Arsenic removal results from these tests are shown
in Table 2-6.
TABLE 2-6
Arsenic Removal by UF at Pilot-Scale
Membrane
Desal GM2540F
Desal GM2540F
MWCO
8,000
8,000
Water Type
High DOC GW
Low DOC GW
Finished SW
As Species
Total As
Total As
V
m
As Rejection
70%
30%
47%
10%
As seen in Table 2-6, arsenic removal varied with DOC levels, being much higher in the high
DOC groundwater (70%) than in the low DOC groundwater (30%). The authors postulated that this
2-33
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difference was due to a reduction in electrostatic forces caused by adsorption of NOM to the
membrane surface. Adsorption of NOM would reduce the surface charge of the membrane and
would, in effect, increase the repulsion towards negatively charged arsenic compounds. Increases
in the apparent size of the arsenic molecules through "bridging" with humic substances was ruled
out since a concurrent increase in UV2S4 removal was not seen. In contrast to DOC levels, changes
in flux and recovery did not seem to impact the arsenic rejection rate. As shown in Table 2-6, testing
on the finished surface water showed fairly effective removal of As(V), but unimpressive As(III)
removal. Considering the MWCO, however, these removals were expected.
2.5.6 Nanofiltration
Nanofiltration membranes are capable of removing significant portions of the dissolved
arsenic compounds in natural waters due to their small pore size. NF will primarily remove divalent
ions (e.g., Ca, Mg), but not monovalent salts (e.g., Na, Cl). Through size exclusion, NF can remove
both dissolved As(V) and As(III). This makes NF a reliable arsenic removal process for
groundwater which contains up to 90% dissolved arsenic (AWWARF, 1998). The small pore size,
however, makes NF membranes more prone to fouling than UF or MF membranes. The application
of NF for surface water treatment is typically not accomplished without extensive pretreatment for
particle removal and possibly pretreatment for dissolved constituents to prevent fouling.
Several NF studies for have been undertaken, and the results show that NF processes are
effective for the removal of arsenic. Removal however depends on operating parameters, membrane
properties, and arsenic speciation. AWWARF (1998) performed NF bench-scale studies for arsenic
removal on spiked deionized water and on a lake water. Single element and flat sheet testing were
performed on a negatively charged NF membrane for a lake water and a spiked, deionized water.
Results are shown in Table 2-7.
As seen in Table 2-7, As(III) removal was low at only 12 percent. However, As(V) rejection
for the negatively charged membrane was liigh at 89 and 85 percent for the lake water and deionized
water, respectively. Flat sheet testing produced a comparable As(V) rejection of 90 percent.
2-34
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TABLE 2-7
As(V) and As(III) Removal by NF Membranes
Membrane Type
Single Element
NF 45-2540
NF 45-2540
NF 45-2540
Flat Sheet
NF 45-2540
MWCO
300
300
300
300
Membrane
Charge
(-)
(-)
(-)
(-)
Water
Type
DI
Lake
DI
DI
Species
V
V
m
V
PH
•
6.7
6.9
6.9
NA
Total Arsenic
Rejection
(%)
85
89
12
90
NA: Not Available
AWWARF also performed several single element and array NF pilot-scale tests. Two of
these tests were conducted on groundwaters, one high in DOC (11 mg/L) and one low in DOC (1
mg/L). Another test was performed on spiked, high-DOC groundwater. One other test was
performed on spiked, finished surface water. These tests are summarized in Table 2-8.
As shown in Table 2-8, during the single element tests on the groundwaters the membranes
demonstrated substantial arsenic removal. Removal in the low DOC water, however, was only 60
percent compared to over 80 percent in the high DOC water. As discussed in Section 2.5.5, this was
presumably due to changes in electrostatic repulsion at the membrane surface through NOM
adsorption. As in the UF pilot study, NF arsenic rejection rate did not seem to be affected by
changes in flux or recovery.
Single element tests performed on the spiked, finished surface water showed substantial
As(V) rejection (>95 percent). As(III) rejection, however, was reduced with an average for all three
membranes of only 40 percent. The authors point out that these results attest to the influence of
diffusion and electrostatic repulsion on As(III) removal. As(III) is small and can more easily diffuse
through very small NF pores. As(III) is also not as repulsed by surface charge as As(V). Combining
NF with an oxidizing process to convert As(III) to As(V) would probably be the most effective
option for its removal.
2-35
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TABLE 2-8
Arsenic Removal with NF at Pilot-Scale
Membrane
Single Element
Accumem
NF1
NF2
NFS
Array
Accumem
MWCO
400
NA
NA
NA
400
Water Type
High DOC GW
Low DOC GW
Finished SW
Finished SW
Finished SW
High DOC GW
Charge
(-)
(-)
NA
NA
NA
(-)
As
Species
,
Total As
Total As
V
III
V
m
V
m
Total As
As
Rejection
80%
60%
>95%
52%
>95%
20%
>95%
30%
75% (initial)
3- 16% (final)
NA: Not Available
The array test results, as shown in Table 2-8, were somewhat surprising. Arsenic rejection
rate declined over time. Rejection at the beginning of the test was approximately 75 percent but
proceeded to decline to 11 percent by day 60. Rejection stayed between 3 percent and 16 percent
for the remainder of the 80-day period. This was surprising given the fact that the membrane
showed high arsenic rejection in single-element tests. Samples taken throughout the array indicate
that a speciation change from As(V) to As(III) was taking place within the filter. Since As(III) is
more difficult to remove than As(V), overall arsenic removal dropped. This decrease in rejection
over time suggests that a negatively charged membrane could not ke<;p high As (V) rejection rates
for long durations without maintaining arsenic in the As(V) form. Additional long-term testing is
needed to verify these results for other membranes and situations. If speciation changes are
influential for arsenic removal, keeping the membrane surface in an oxidized state may be an option.
A NF pilot-scale study to determine arsenic removals with NF membranes was conducted
in Tarrytown, NY (Malcolm Pimie 1992). Two NF membranes were tested: (1) NF70 manufactured
2-36
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by Dow Chemical Company (FilmTec), and (2) TFCS manufactured by UOP Fluid Systems. The
NF membranes were operated at a flux varying between 17 and 21 gfd and at a recovery of 15
percent. Feed water conductivity varied from 460 to 950 uS, pH ranged from 7.7 to 8.3, and feed
water arsenic ranged from 0.038 - 0.154 mg/L. A second feed solution was mixed that had
approximately twice the TDS and arsenic levels as found in the original test solution to simulate
arsenic rejections by the last element in an NF membrane system operating at 50 percent recovery.
Arsenic rejection was very high with only one of eight permeate samples from the NF membranes
exceeding the detection limit with a level of 0.0025 mg/L, corresponding to 95% rejection.
2.5.7 Reverse Osmosis
RO is the oldest membrane technology, traditionally used for the desalination of brackish
water and sea water. RO produces nearly pure water by maintaining a pressure gradient across the
membrane greater than the osmotic pressure of the feed water. Osmotic pressure becomes great in
RO systems compared to other membrane processes due to the concentration of salts on the feed side
of the membrane. The majority of the feed water passes through the membrane, however, the rest
is discharged along with the rejected salts as a concentrated stream. Discharge concentrate can be
substantial, between 10 and 50 percent of the influent flow depending on influent water quality and
membrane properties.
RO performance is adversely affected by the presence of turbidity, iron, manganese, silica,
scale-producing compounds, and other constituents. Like NF, RO requires extensive pretreatment
for particle removal and often pretreatment for dissolved constituents. RO often requires
pretreatment even for high quality source waters. RO has sometimes been used as a polishing step
for already treated drinking water. Pretreatment can make RO processes costly. Treated waters from
RO systems typically have extremely high quality, however, and blending of treated water and raw
water can be used to produce a finished water of acceptable quality. This may reduce cost to some
extent.
RO is an effective arsenic removal technology proven through several bench- and pilot-scale
studies, and is very effective in removing dissolved constituents. Since the arsenic found in
groundwater is typically 80 to 90 percent dissolved, RO is a suitable technology for arsenic removal
in groundwater. Several previous RO bench-scale and pilot-scale studies for arsenic removal are
2-37
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summarized in Table 2-9. These studies indicate that RO can be an effective process for arsenic
removal, however, membrane type and operating conditions will affect removal and must be chosen
appropriately. As with other processes, RO removes As(V) to a greater degree than As(III), so
maintaining oxidation conditions may be important to the process.
AWWARF (1998) performed bench- and pilot-scale RO testing. Short-term, single element
testing and flat sheet testing were performed for a DK2540F RO membrane manufactured by
DESAL on a lake water and on spiked deionized water. Results from tiiis testing are shown in Table
2-10.
TABLE 2-9
Summary of Arsenic Removal with RO
Location
Eugene, OR
Eugene, OR
Fairbanks, AL
San Ysidro, NM
San Ysidro, NM
San Ysidro, NM
Tarrytown, NY
Tarrytown, NY
Charlotte Harbor,
FL
Cincinnati, OH
Hudson, NH
Type
POU
POU
POU
Pilot
(hollow fiber, cellulose
acetate)
Pilot
(hollow fiber,
polyamide)
POU
Pilot
(FilmTec BW30,
Hydranautics NCM1,
Fluid Systems TCFL)
POU
POU
(several membrane
types)
POU
POU
Operating Parameters
• 3-5 gpd
• 90% recovery
• 20-100psi
• 3-5 gpd
• 67% recovery
• 195 psi
• low-pressure (< 100
psi)
50% recovery
pH adjustment to 6.3
antiscalent addition
50% recovery
pH adjustment to 6.3
antiscalent addition
• 10-1 5% recovery
• 15 gfd
• 10% recovery
NA
• 1000 gpd
• 10-60% recovery
NA
NA
As Removal
50%
below MDL
50%
93-99%
99%
91%
below MDL
86%
As(V) 96-99%
As(III) 46-84%
As(IIl) 73%
40%
Reference
Fox, 1989
FoxandSorg, 1987
Fox, 1989
Fox and Sorg, 1987
Fox, 1989
FoxandSorg, 1987
Clifford and Lin, 1991
Clifford and Lin, 1991
Fox, 1989
FoxandSorg, 1987
Malcolm Pirnie, 1992
Rogers, 1989
Huxstep, 1987
Fox and Sorg, 1987
USEPA, 1982
NA: Not Available
2-38
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TABLE 2-10
Arsenic Removal with RO at Bench-Scale
Membrane
Single Element
DK2540F
DK2540F
DK2540F
DK2540F
Flat Sheet
DK2540F
MWCO
180
180
180
180
180
Source Water
Deionized
Lake Water
Deionized
Lake Water
Deionized
Speciation
V
V
m
m
V
pH
'6.8
6.9
6.8
6.8
As Rejection
96%
96%
5%
5%
88%
These results indicate very high rejection for As(V) but very low rejection for As(III) at
neutral pH. Again, this points to the fact that oxidation conditions would be desirable, and that
surface charge/electrostatic repulsion probably plays a role in arsenic rejection. Also, flat sheet
testing produced a rejection rate comparable, and slightly conservative, to the single element
rejection rate.
Several RO pilot-scale tests were also performed (AWWARF, 1998). Two tests were
performed on high- and low-DOC groundwaters. Another set of tests was performed on spiked,
finished surface water. The results from these pilot tests are summarized in Table 2-11.
Table 2-11 shows substantial rejection for both the low- and high-DOC waters. Rejection
was only slightly higher with the high DOC water. As with UF and NF, flux and recovery changes
did not seem to affect arsenic rejection. Results for the four membranes tested on spiked finished
water also showed substantial removal. For all membranes during this test, As(V) exceeded 95
percent, however, As(III) rejection averaged only 74 percent.
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TABLE 2-11
Arsenic Removal with RO at Pilot-Scale
Membrane
Single Element
TFCL-HR
R01
RO2
RO3
RO4
Water Type
High DOC GW
Low DOC GW
Finished SW
Finished SW
Finished SW
Finished S W
Charge
(-)
(-)
NA
NA
NA
NA
As
Species
Total As
Total As
V
in
V
m
V
in
V
m
As
Rejection
>90%
>80%
>95%
60%
>95%
75%
>95%
68%
>95%
85%
NA: Not Available
Overall, RO is capable of achieving finished water arsenic concentrations below 0.002 mg/L
when arsenic is present as As(V). As(III) rejection is not as significant, however, conversion to
As(V) can be achieved with pre-oxidation.
2.5.8 Electrodialysis Reversal
Electrodialysis (ED) is a process in which ions are transferred through membranes that are
selectively permeable towards cations or anions under the influence of direct electric current. The
separation mechanism is actually an ion exchange process. The ions travel from a lesser to a higher
concentrated solution. In this process, the membranes are arranged in an array or stack placed
between opposite electrodes, with alternating cation and anion exchange membranes. The mobility
of the cations or anions is restricted to the direction of the attracting electrodes, and this results in
alternating sets of compartments containing water with low and high concentrations of the ions. The
electrodialysis reversal (EDR) process is an ED process with periodic reversal of the direction of
travel of the ions caused by reversing the polarity of the electrodes. The advantage of polarity
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reversal is the decreased potential for fouling of the membranes, which also minimizes the
pretreatment requirements of this process.
EDR is designed specifically for each application based on the desired quantity and quality
of product water. Equipment at an EDR plant, besides the stack itself, includes feedwater pumps,
recycle pumps, valving, stream switching, product water diversion, pressure regulation, and
electrode stream control. EDR systems are fully automated and require little operator attention, with
the exception of data collection and routine maintenance. Routine maintenance consists of changing
cartridge filters, calibrating and maintaining instruments, replacing membranes, maintaining pumps
and valves, and replacing electrodes. EDR systems are also attractive since they do not require
chemical addition (EPA, 1994). EDR systems, however, are typically more expensive than NF and
RO systems (EPA, 1994). EDR systems are often used in treating brackish water to make it suitable
for drinking, hi terms of effluent water quality, EDR has been compared to RO (AWWARF, 1996;
Robinson, et al., 1998). EDR processes have also been applied in the industry for wastewater
recovery.
EDR can achieve high removals of TDS from water and typically operates at a recovery of
70 to 80 percent (Kempic, 1994a). Very few studies have been conducted to exclusively evaluate
this process for the removal of arsenic. One of the studies was conducted using EDR to treat water
from San Ysidro, New Mexico, which was a site for several other arsenic removal studies (Clifford
and Lin, 1985). Studies by a leading manufacturer of EDR equipment also provide data on arsenic
removal (Ionics Inc., 1989-1990). These are discussed below.
In the San Ysidro EDR study, a recovery of 85 percent was achieved by using an internal
brine recycle system. Pretreatment for the unit consisted of a standard 10-micron cartridge filter and
a granular activated carbon (GAC) column that were part of the system provided by the manufactur-
er. The unit was tested for two different waters, a city water that contained a mixture of As(III) and
As(V), and a groundwater that contained mostly As(III). The well water contained 0.188 mg/L of
arsenic. The groundwater was nearly all As(III). Arsenic removals by EDR were low, at only 28
percent, and the effluent concentrations were high at 0.136 mg/L.
The city water quality is shown in Table 2-12. The unit was run for 5 days with a recovery
of 81 percent. The overall removal of arsenic was estimated at 73 percent. Approximately 60
percent of the As(III) was removed, which was higher than expected.
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Another mobile unit equipped with RO, ED, and EDR systems, along with the necessary
pretreatment and post-treatment equipment, was used to test waters from eight New Mexico
communities (New Mexico State University, 1979). In one of the studies conducted at Bluewater,
New Mexico, EDR brought the level of areenic in the treated water down to 0.003 mg/L from the
influent level of 0.021 mg/L. This corresponds to a removal of approjdmately 86 percent. The feed
water to the EDR unit was drawn from a point before chlorination of the community water supply.
The test flow rate was 4.8 gpm, and 80 percent recovery was obtained. Raw water quality for the
community water is shown in Table 2-13.
TABLE 2-12
Influent Water Quality for San Ysidro EDR Study
Parameter
pH
TDS
As(total)
Fluoride
Sulfate
Bicarbonate
Chloride
Concentration (mg/L)
7.1 (units)
810
0.085
2.4
36
552
142
2-42
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TABLE 2-13
Raw Water Quality for Bluewater EDR Study
Parameter
pH (units)
TDS
Na
Sulfate
Silica
Chloride
Concentration
(mg/L)
7.1
908
78
398
16
52
In another study, process water from in-situ mining was treated using a 30,000-gpd EDR unit
(Garling,1981). The unit removed about 59 percent of the 0.022 mg/L arsenic in the feedwater
operating at a recovery of approximately 81 percent.
2.6 ALTERNATIVE TECHNOLOGIES
2.6.1 Oxidation Filtration
Oxidation filtration technologies may be effective arsenic removal technologies. Research
of oxidation filtration technologies has primarily focused on greensand filtration. As a result, this
discussion focuses on the effectiveness of greensand filtration as an arsenic removal technololgy.
Substantial arsenic removal has been seen using greensand filtration (Subramanian, et al.,
1997). The active material in "greensand" is glauconite, a green, iron-rich, clay-like mineral that has
ion exchange properties. Glauconite often occurs in nature as small pellets mixed with other sand
particles, giving a green color to the sand. The glauconite sand is treated with KMnO4 until the sand
grains are coated with a layer of manganese oxides, particularly manganese dioxide. The principle
behind this arsenic removal treatment is multi-faceted and includes oxidation, ion exchange, and
adsorption. Arsenic compounds displace species from the manganese oxide (presumably OH" and
H2O), becoming bound to the greensand surface - in effect an exchange of ions. The oxidative nature
of the manganese surface converts As(III) to As(V) and As(V) is adsorbed to the surface. As a result
2-43
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of the transfer of electrons and adsorption of As(V), reduced manganese (Mnll) is released from the
surface.
The effectiveness of greensand filtration for arsenic removal is dependent on the influent
water quality. Subramam'an et al. (1997) showed a strong correlation between influent Fe(II)
concentration and arsenic percent removal. Removal increased from 41 percent to more than 80
percent as the Fe/As ratio increased from 0 to 20 when treating a tap water with a spiked As(III)
concentration of 200 ug/L. The tap water contained 366 mg/L sulfate and 321 mg/L TDS; neither
constituent seemed to affect arsenic removal. The authors also point out that the influent Mn(IV)
concentration may play an important role. Divalent ions, such as calcium, can also compete with
arsenic for adsorption sites. Water quality would need to be carefully evaluated for applicability for
treatment using greensand. Other researchers have also reported substantial arsenic removal using
this technology, including arsenic removals of greater than 90 percent for treatment of groundwater
(Subramanian, et al., 1997).
As with other treatment media, greensand must be regenerated when its oxidative and
adsorptive capacity has been exhausted. Greensand filters are regenerated using a solution of excess
potassium permanganate (KMnO4). Like other treatment media, the regeneration frequency will
depend on the influent water quality in terms of constituents which will degrade the filter capacity.
Regenerant disposal for greensand filtration has not been addressed in previous research.
2.6.2 Sulfur-Modified Iron
A patented Sulfur-Modified Iron (SMI) process for arsenic; removal has recently been
developed (Hydrometrics, 1997 and 1998). The process consists of lihree components: (1) finely-
divided metallic iron; (2) powdered elemental sulfur, or other sulfur compounds; and (3) an
oxidizing agent. The powdered iron, powdered sulfur, and the oxidizing agent (H2O2 in preliminary
tests) are thoroughly mixed and then added to the water to be treated. The oxidizing agent serves
to convert As(III) to As(V). The solution is then mixed and settled.
Using the SMI process on several water types, high adsorptive capacities were obtained with
final arsenic concentration of 0.050 mg/L. Arsenic removal was influenced by pH. Approximately
20 mg As per gram of iron was removed at pH 8, and 50 mg As per gram of iron was removed at
pH 7. Arsenic removal seems to be very dependent on the iron to arsenic ratio.
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Packed bed column tests demonstrated significant arsenic removal at residence times of 5
to 15 minutes. Significant removal of both arsenate and arsenite was measured. The highest
adsorption capacity measured was 11 mg As removed per gram of iron. Flow distribution problems
were evident, as several columns became partially plugged and better arsenic removal was observed
with reduced flow rates.
Spent media from the column tests were classified as nonhazardous waste. Projected
operating costs for SMI, when the process is operated below a pH of 8, are much lower than
alternative arsenic removal technologies such as ferric chloride addition, reverse osmosis, and
activated alumina. Cost savings would increase proportionally with increased flow rates and
increased arsenic concentrations.
Possible treatment systems using SMI include continuous stirred tank reactors, packed bed
reactors, fluidized bed reactors, and passive in situ reactors. Packed bed and fluidized bed reactors
appear to be the most promising for successful arsenic removal in pilot-scale and full-scale treatment
systems based on present knowledge of the SMI process.
2.6.3 Granular Ferric Hydroxide
A new removal technique for arsenate, which has recently been developed at the Technical
University of Berlin (Germany), Department of Water Quality Control, is adsorption on granular
ferric hydroxide (GFH) in fixed bed reactors. This technique combines the advantages of the
coagulation-filtration process, efficiency and small residual mass, with the fixed bed adsorption on
activated alumina, and simple processing.
Driehaus et al. (1998) reported that the application of GFH in test adsorbers showed a high
treatment capacity of 30,000 to 40,000 bed volumes with an effluent arsenate concentration never
exceeding 10 ug/L. The typical residual mass was in the range of 5-25 g/m3 treated water. The
residue was a solid with an arsenate content of 1-10 g/kg. Table 2-14 summarizes the data of the
adsorption tests.
2-45
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Table 2-14
Adsorption Tests on GFH
Units
Testl
Test 2
TestS
Test 4
Raw Water Parameters
pH
Arsenate Concentration
Phosphate Concentration
Conductivity
Adsorption Capacity for Arsenate
Hg/L
"g/L
uS/cm
g/kg
7.8
100-800
0.70
780
8.5
7.8
21
0.22
480
4.5
8.2
16
0.15
200
3.2
7.6
15-20
0.30
460
N/D
Adsorber
Bed Height
Filter Rate
Treatment Capacity
Maximum Effluent Concentration
Arsenate Content of GFH
Mass of Spent GFH (dry weight)
m
m/h
BV
Hg/L
g/kg
g/m3
0.24
6-10
34,000
10
8.5
20.5
0.16
7.6
37,000
10
1.4
12
0.15
5.7
32,000
10
0.8
18
0.82
15
85,000
7
1.7
8.6
N/D: not determined
The competition of sulfate on arsenate adsorption was not very strong. Phosphate, however,
competed strongly with arsenate, which reduced arsenate removal with GFH. Arsenate adsorption
decreases with pH, which is typical for anion adsorption. At high pH values GFH out-performs
alumina. Below a pH of 7.6 the performance is comparable. The most significant weakness,
however, appears to be cost. Currently, GFH media costs approximately $4,000 per ton. The effect
on total O&M costs should be evaluated. It is possible that if a GFH bed can be used several times
longer than an alumina bed, for example, die overall effect may be minimal.
A treatment for leaching arsenic and the regeneration of GFH seems possible, but it leads to
an alkaline solution with arsenate and requires a further treatment to obtain a solid waste. Thus, a
direct deposition of spent GFH as hazardous waste should be favored.
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2.6.4 Iron Filings
Iron filings and sand may be used to reduce inorganic arsenic species to iron co-precipitates,
mixed precipitates and, in conjunction with sulfates, to arsenopyrites. This type of process is
essentially a filter technology, much like greensand filtration, wherein the source water is filtered
through a bed of sand and iron filings. Unlike some technologies, ion exchange for example, sulfate
is actually introduced in this process to encourage arsenopyrite precipitation.
This arsenic removal method was originally developed as a batch arsenic remediation
technology. It appears to be quite effective in this use. Bench-scale tests indicate an average
removal efficiency of 81% with much higher removals at lower influent concentrations. This
method was tested to arsenic levels of 20,000 ppb, and at 2000 ppb consistently reduced arsenic
levels to less than 50 ppb (the current MCL). While it is quite effective in this capacity, its use as
a drinking water treatment technology appears to be limited. In batch tests a residence time of
approximately seven days was required to reach the desired arsenic removal. In flowing conditions,
even though removals averaged 81% and reached greater than 95% at 2000 ppb arsenic, there is no
indication that this technology can reduce arsenic levels below approximately 25 ppb, and there are
no data to indicate how the technology performs at normal source water arsenic levels. This
technology needs to be further evaluated before it can be recommended as an approved arsenic
removal technology for drinking water.
2.6.5 Photo-Oxidation
Researchers at the Australian Nuclear Science and Technology Organisation (ANSTO) have
found that in the presence of light and naturally occurring light-absorbing materials, the oxidation
rate of As(III) by oxygen can be increased ten-thousandfold (Cooperative Research Centres for
Waste Management and Pollution Control Limited, 1999). The oxidized arsenic, now As(V), can
then be effectively removed by co-precipitation.
ANSTO evaluated both UV lamp reactors and sunlight-assisted-photo-oxidation using acidic,
metal-bearing water from an abandoned gold, silver, and lead mine. Air sparging was required for
sunlight-assisted oxidation due to the high initial As(III) concentration (12 mg/L). Tests
demonstrated that near complete oxidation of As(III) can be achieved using the photochemical
process. Analysis of process waters showed 97% of the arsenic in the process stream was present
2-47
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as As(V). Researchers also concluded that As(III) was preferentially oxidized in the presence of
excess dissolved Fe(II) (22:1 iron to arsenic mole ratio). This is a contrast to conventional plants
where dissolved Fe(II) represents an extra chemical oxidant demand which has to be satisfied during
oxidation of As(III) (CRC-WMPC, 1999).
Photo-oxidation of the mine water followed by co-precipitation was able to reduce arsenic
concentrations to as low as 17 yug/L, which meets the current MCL for arsenic. Initial total arsenic
concentrations were unknown, though the As(III) concentration was given as approximately 12
mg/L, which is considerably higher than typical raw water arsenic concentrations. ANSTO reported
residuals from this process are environmentally stable and passed the Toxicity Characteristic
Leaching Procedure (TCLP) test necessary to declare waste non-hazardous and suitable for landfill
disposal.
Based upon the removals achieved and residuals characteristics, it is expected that photo-
oxidation followed by co-precipitation would be an effective arsenic removal technology. However,
this technology is still largely experimental and should be further evaluated before recommendation
as an approved arsenic removal technology for drinking water.
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3.0 TECHNOLOGY COSTS
3.1 INTRODUCTION
This chapter presents estimated capital and operations and maintenance (O&M) expenditures
for the following arsenic removal technologies and unit processes:
• Pre-oxidation technologies, including chlorination and potassium permanganate;
• Precipitative processes, including coagulation/filtration, direct filtration, coagulation
assisted microfiltration, enhanced coagulation, lime softening, and enhanced lime
softening;
• Adsorption processes, including activated alumina;
• Ion exchange processes, specifically anion exchange;
• Separation processes, including ultrafiltration, nanofiltration, and reverse osmosis.
Each section includes a brief technology description, design criteria, and capital and O&M
cost curves for systems with design flows ranging from 0.01 to 430 mgd.
3.2 BASIS FOR COST ESTIMATES
3.2.1 Cost Modeling
Three cost models were used in cost development: the Very Small Systems Best Available
Technology Cost Document (Malcolm Pirnie, 1993), hereafter referred to as the VSS model; the
Water Model (Culp/Wesner/Culp, 1984); and the WAV Cost Model (CulpAVesner/Culp, 1994).
Curve fitting analysis was conducted on the modeled cost estimates including the utilization of
transition flow regions to provide better estimates within the breakpoints between models.
3-1
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The following flow ranges have been established for each model and transition flow region:
VSS - 0.015 to 0.100 mgd
Transition 1 - 0.100 to 0.270 mgd
Water Model - 0.27 to 1.00 mgd
Transition 2 - 1 to 10 mgd
WAV Cost Model - 10 to 200 mgd
Some processes (e.g., activated alumina and ion exchange) have slightly different ranges due
to discrepancies between the models. Membrane processes have diffident ranges since these costs
were generally developed without use of the models. All three models require flow to calculate
capital and operation and maintenance (O&M) costs. In addition, the Water and WAV Cost models
require several user-specified variables to generate direct capital cost. These additional user inputs
include design factors, cost indices (Table 3-7), and other various unit costs (Tables 3-8 and 3-9).
3.2.2 Technology Design Panel Recommendations
Since the 1986 Safe Drinking Water Act (SDWA) reauthorization, EPA has relied mainly
on the previously mentioned unit cost models to estimate compliance costs for drinking water
regulations. Following the reauthorization of the SDWA in 1996, EPA has critically evaluated its
tools for estimating the costs and benefits of drinking water regulations. As part of this evaluation,
EPA solicited technical input from national drinking water experts at the Denver Technology
Workshop (which was sponsored by EPA and held November 6 and 7,, 1997) to improve the quality
of its compliance cost estimating process for various drinking water treatment technologies. The
Technology Design Panel (TOP) formed at the workshop for this purpose recommended several
modifications to existing cost models to improve the accuracy of EPA's compliance cost estimates.
The TDP developed guidelines for estimating capital costs using the three cost models. The
guidelines are discussed in greater detail in Guide for Implementing Phase I Water Treatment
Upgrade (EPA, 1998a) and Water Treatment Costs Development (Phase I): Road Map to Cost
Comparisons (EPA, 1998b).
Total capital costs consist of three elements: process, construction, and engineering costs.
Process costs include manufactured equipment, concrete, steel, electrical and instrumentation, and
3-2
-------
pipes and valves. Construction costs include sitework and excavation, subsurface considerations,
standby power, land, contingencies, and interest during construction. Engineering costs include
general contractor overhead and profit, engineering fees, and legal, fiscal, and administrative fees
(including permitting). Housing costs are specifically excluded from each of these cost category
designations. Housing costs are included in the estimates presented in this chapter, but are added
to the total capital cost after application of the TDP cost approach.
The TDP recommended that total capital cost estimates be generated based solely upon
process costs. That is, the models can be used to estimate total capital costs, but process costs are
then generated using the capital cost breakdowns presented in Appendices A through C, and
applying an appropriate factor for construction and engineering costs. These factors are based upon
system size and are presented in Table 3-1.
Table 3-1
TDP Capital Cost Factors
System Size
Very Small
Small
Large
Process Cost
Factor
(Percent of Total)
1.00 (40%)
1.00(40%)
1.00(30%)
Construction Cost
Factor
(Percent of Total)
1.00 (40%)
1.00(40%)
1.33 (40%)
Engineering Cost
Factor
(Percent of Total)
0.50 (20%)
0.50 (20%)
1.00(30%)
Total Cost
Factor1
(Percent of Total)
2.50 (100%)
2.50 (100%)
3.33 (100%)
1 - This factor can be multiplied by the process cost to obtain the total capital cost excluding housing. Housing costs are added to the total cost.
Table 3-2 presents a sample capital cost breakdown for the VSS model membrane equations.
The table also lists the capital costs assumptions associated with the VSS model. Capital cost
breakdowns for all technologies costed using the VSS model are presented in Appendix A.
The Water and WAV Cost assumptions for capital cost components vary by design and
average flow. Supporting documentation was used to develop capital cost breakdown summaries
for the Water and WAV Cost models. Estimation of Small System Water Treatment Costs
(CulpAVesner/ Gulp, 1984) and Estimating Treatment Costs, Volume 2: Cost Curves Applicable to
1 to 200 mgd Treatment Plants (CulpAVesner/Culp, 1979) were used for the Water and WAV Cost
models, respectively. These documents present the design assumptions used in developing the cost
models, as well as associated costs. The percent of total cost for each component cost was calculated
3-3
-------
for each design condition. These percentages were averaged to arrive at a universal capital cost
breakdown which could be applied for developing the Phase I capital costs. Tables 3-3 through 3-6
demonstrate the methodology described here.
Table 3-2
VSS Capital Cost Breakdown for Membrane Processes
(Including Microfiltration and Ultrafiltration)
Cost Component
Manufactured Equipment
Installation
Sitework and Interface Piping
Standby Power
General Contractor Overhead & Profit
Legal, Fiscal and Administrative Fees
Engineering
Miscellaneous and Contingencies
TOTAL
Model
Assumption
100%
25%
6%
5%
12%
3%
10%
0%
Cost Factor
1.000
0.2500
0.0750
0.0625
0.1665
0.0416
0.1596
0.000
1.7552
Percent of
Total Capital
56.97%
14.24%
4.27%
3.56%
9.49%
2.37%
9.09%
0.00%
100.00%
Capital Cost
Category
P
c
c
c
~e
e
e
c
p = process, c = construction, e = engineering
Output from the Water and WAV Cost models includes construction costs and additional
capital costs, which together make up the total capital cost. Additional capital costs include sitework
and interface piping, standby power, overhead and profit, engineering, legal, fiscal, and
administrative fees. There are no process costs associated with the additional capital costs. As a
result, cost breakdowns need only consider the construction cost output from these two models.
Tables 3-4 and 3-6 present sample capital cost breakdowns for the Water and WAV Cost models,
respectively. Capital cost breakdowns for each technology and unit process are presented in
Appendices A through C for the VSS, Water, and WAV Cost models, respectively.
3-4
-------
Table 3-3
Water Model Capital Cost Breakdown for
Package Conventional Treatment (Coagulation/Filtration)
Cost Component
Excavation and Sitework
Manufactured Equipment
Concrete
Labor
Pipes and Valves
Electrical
Housing*
Subtotal
Contingencies
Total
Filter Area (ft2)
2
$3,500
$31,000
$1,000
$9,900
$4,200
$3,200
$18,600
$71,400
$10,700
$82,100
12
$3,500
$44,900
$1,000
$14,700
$8,300
$4,500
$18,600
$95,500
$14,300
$109,800
20
$4,700
$53,500
$1,500
$17,500
$10,400
$5,300
$23,400
$116,300
$17,400
$133,700
40
$5,800
$111,300
$4,500
$36,400
$20,900
$11,100
$45,000
$235,000
$35,300
$270,300
112
$7,000
$176,600
$5,700
$57,800
$29,200
$17,600
$47,500
$341,400
$51,200
$392,600
150
$9,300
$190,500
$6,800
$62,400
$41,700
$19,000
$52,500
$382,200
$57,300
$439,500
Capital
Cost
Category
c
P
P
c
P
P
c
' ''- '"• •:''• 't, ;'•'"•
e
',' '„ , •:<,''", ':',if:
: ,-.W". ,—. .\~4
'Housing costs are added to the total capital cost after application of the TDP cost approach
Table 3-4
Water Model Capital Cost Breakdown by Percentage for
Package Conventional Treatment (Coagulation/Filtration)
Cost Component
Excavation and Sitework
Manufactured Equipment
Concrete
Labor
Pipes and Valves
Electrical
Housing*
Contingencies
Total
Filter Area (ft2)
2
4.26%
37.76%
1.22%
12.06%
5.12%
3.90%
22.66%
13.03%
100.00%
12
3.19%
40.89%
0.91%
13.39%
7.56%
4.10%
16.94%
13.02%
100.00%
20
3.52%
40.01%
1.12%
13.09%
7.78%
3.96%
17.50%
13.01%
100.00%
40
2.15%
41.18%
1.66%
13.47%
7.73%
4.11%
16.65%
13.06%
100.00%
112
1.78%
44.98%
1.45%
14.72%
7.44%
4.48%
12.10%
13.04%
100.00%
150
2.12%
43.34%
1.55%
14.20%
9.49%
4.32%
11.95%
13.04%
100.00%
Average
Percent
2.84%
41.36%
1.32%
13.49%
7.52%
4.15%
16.30%
13.03%
100.00%
•Housing costs are added to the total capital cost after application of the TDP cost approach
3-5
-------
Table 3-5
WAV Cost Model Capital Cost Breakdown for Sedimentation Basins
Cost Component
Excavation and Sitework
Manufactured Equipment
Concrete
Steel
Labor
Pipes and Valves
Electrical
Subtotal
Contingencies
Total
Area (A = ft2) and Length x Width (LW == ft x ft)
A=240
LW =
30x8
$1,060
$8,540
$2,970
$6,400
$6,220
$6,960
$1,510
$33,660
$5,050
$38,710
A=600
LW=60xlO
$2,000
$12,080
$5,490
$13,110
$11,260
$7,400
$1,760
$53,100
$7,970
$61,070
A=1260
LW=90xl4
$3,060
$24,470
$84,430
$19,440
$17,320
$9,100
$1,860
$83,680
$12,550
$96,230
A=2240
LW=140xl6
$4,680
$32,020
$12,820
$32,620
$26,390
$12,500
$2,020
$123,050
$18,460
$141,510
A=S600
LW=:tOOxl8
$6,670
$53,110
$19,190
$51,250
$37,570
$16,100
$2,110
$190,000
$27,750
$212,750
A=4800
LW=240x20
$8,090
$63,440
$22,070
$39,680
$45,300
$21,450
$2,400
$232,430
$34,860
$267,290
Capital
Cost
Category
c
P
P
P
c
P
P
< s,^ ^>-\ , > :
yl' * *' ,f.
e
,>y *§'s "-.if "v!'
Table 3-6
WAV Cost Model Capital Cost Breakdown by Percentage for Sedimentation Basins
Cost Component
Excavation and Sitework
Manufactured Equipment
Concrete
Steel
Labor
Pipes and Valves
Electrical
Contingencies
Total
Area (A = ft2) and Length x Width (LW = ft x ft)
A=240
LW =
30x8
2.74%
22.06%
7.67%
16.53%
16.07%
17.98%
3.90%
13.05%
100.00%
A=600
LW=60xlO
3.27%
19.78%
8.99%
21.47%
18.44%
12.12%
2.88%
13.05%
100.00%
A=1260
LW=90xl4
3.18%
25.43%
8.76%
20.20%
18.00%
9.46%
1.93%
13.04%
100.00%
A=2240
LW=140xl6
3.31%
22.63%
9.06%
23.05%
18.65%
8.83%
1.43%
13.05%
100.00%
A-3600
LW=-200xl8
3.14%
27.96%
8.55%
24.09%
17.66%
7.57%
0.99%
13.04%
100.00%
A=4800
LW=240x20
3.03%
23.73%
8.26%
26.07%
16.95%
8.02%
0.90%
13.04%
100.00%
Average
Percent
3.11%
23.10%
8.55%
21.90%
17.63%
10.66%
2.01%
13.04%
100.00%
3-6
-------
3.2.3 Implementing TDP Recommended Costing Upgrades
The capital cost breakdowns presented above and in the appendices of this document can be
used to estimate the modified capital cost, i.e., the capital cost estimate developed using the TDP
recommendations. The following sections briefly demonstrate how the capital cost breakdowns are
applied, and modified capital cost estimates are generated.
3.2.3.1 VSS Model
1. The VSS model presents capital and O&M costs as functions of design and average flow,
respectively. Accordingly, the capital cost equation for package microfiltration units is:
CAP = 0.86[DES]+41.1
Where: CAP = Total Capital Cost, $ 1,000s
DBS = Design Treated Flow, kgpd
2. Thus, for a 0.024 mgd (24 kgpd) plant the capital cost is:
CAP = 0.86[24] + 41.1
CAP = 61.74 or $61,740
3. The VSS model equations produce estimates in 1993 dollars. To escalate to September
1998, multiply the equation-generated capital cost by the ratio of the Engineering News
Record (ENR) Building Cost Index for September 1998 to the 1993 index value.
$61,740 x (3375/3009) = $69,250
The escalated capital cost for a 0.024 mgd package microfiltration plant is $69,250.
4. Using the capital cost breakdown in Table 3-2, the total process cost is:
$69,250 x 0.5697 = $39,452
5. The modified capital cost can then be calculated using the total cost factor presented in
Table 3-1.
$39,452 x 2.5 = $98,629
Thus, the modified capital cost is $98,629.
3-7
-------
3.2.3.2 Water Model
1. Assume the Water model output for a 0.27 mgd (270,000 gpd) package conventional
treatment (coagulation/flocculation/filtration) plant is $692,066 (escalated to 1998
dollars).
2. Using the capital cost breakdown in Table 3-4, the total process cost is:
$692,066 x (0.4136 + 0.0132 + 0.0752 + 0.0415) = $376,138
3. The modified capital cost can then be calculated using the total cost factor presented in
Table 3-1.
$376,138x2.5 = $940,345
4. This approach must be applied to each unit process (e.g., backwash pumping) separately,
then totaled for the entire treatment process to estimate the modified capital cost.
5. When housing costs are included for a unit process, they are added after multiplication
of the process cost by the TDP total cost factor (2.5 in this; example). Table 3-4 shows
housing is 16.3% of the construction cost. The total capital cost is:
($692,066 x 0.1630) + $940,345 = $1,053,152
3.2.3.3 WAV Cost Model
1. Assume the WAV Cost model output for a 1 mgd (1250 ft2) rectangular sedimentation
basin is $416,574 (escalated to 1998 dollars).
2. Using the capital cost breakdown in Table 3-6, the total process cost is:
$416,574 x (0.2311 + 0.0855 + 0.2190 + 0.1066 + 0.0201) = $275,897
3. The modified capital cost can men be calculated using the total cost factor presented in
Table 3-1.
$275,897x3.33 = $918,737.
4. This approach must be applied to each unit process separately (e.g., acid feed), then
totaled for the entire treatment process to estimate the modified capital cost.
3-8
-------
3.2.4 Cost Indices and Unit Costs
Both the Water Model and the WAV Cost Model require a number of standard indices and
various unit costs from the Bureau of Labor Statistics, the Engineering News Record, and other
referenced sources. The values used in conjunction with the development of cost estimates are
reported in Tables 3-7 through Table 3-9.
Table 3-7
Costs Indices Used in the Water and WAV Cost Models
Description
Concrete Ingredients and Related Products
Electrical Machinery and Products
General Purpose Machinery and Equipment
Metals and Metal Products (Steel)
Miscellaneous General Purpose Equipment
(Pipes &Valves)
PPI Finished Goods Index
ENR Building Cost Index
ENR Skilled Labor
ENR Materials Prices
Index
Reference
BLS 132
BLS117
BLS 1 14
BLS 1017
BLS 1 149
BLS 3000
Numerical
Value1
448.8
281.8
445.1
405.1
521.5
364.0
3375.31
5317.36
2189.24
(1) BLS numerical values were re-based to 1967 base year (see Section 3.2.5)
Table 3-8
Unit and General Cost Assumptions
Electricity'
Diesel Fuel1
Natural Gas1
Labor2
Building Energy Use
Housing Costs
$0.08/kWh
$1.25/gallon
$0.006/scf
Large systems:
Small systems:
$40/hr
$28/hr
102.6kWbyft2/yr
5125/ft2
1 Energy Information Administration.
2 Technical Design Panel (EPA, 1998a)
3-9
-------
TABLE 3-9
Chemical Costs
Chemical
Alum, Dry Stock
Carbon Dioxide, Liquid
Chlorine, 1 ton cylinder
Chlorine, 150 Ib cylinder
Chlorine, Bulk
Ferric Chloride
Hexametaphosphate
Lime, Quick Lime
Phosphoric Acid
Polymer
Potassium Permanganate
Soda Ash
Sodium Hypochlorite, 12%
Sodium Chloride
Sodium Hydroxide, 50% solution
Sulfuric Acid
Cost
$300
$340
$350
$400
$280
$350
$1276
$95
$300
$2.25
$2700
$400
$1100
$99
$371
$116
Units
per ton
per ton
per ton
per ton
per ton
per ton
per ton
per ton
per ton
perlb
per ton
per ton
per ton
per ton
per ton
per ton
This document presents total capital costs and annual O&M costs. Annual O&M costs
include the costs for materials, chemicals, power, and labor. Annualized costs can be determined
using the following equations:
Total annual cost (0/kgal) = Annualized Capital Cost (0/kgal) + O&M Cost (0/kgal)
Where:
Annualized Capital Cost = Capital Cost (S) * Amortization Factor * LOO_#
Average Daily Flow (mgd)*(1000 kgal/mgal)*365 days/year
O&M Cost (0/kgal)
Annual O&M ($) * 100 (eV$)
Average Daily Flow (mgd)*1000 kgal/mgal*365 days/year
3-10
-------
Amortization, or capital recovery, factors for interest rates of 3,7, and 10 percent for 20 years
are reported in Table 3-10. Alternative capital recovery factors can be calculated using the formula
presented below.
Capital Recovery Factor = i(l + i)N / (1 + i)N -1
Where: i = interest rate
N = number of years
Table 3-10
Amortization Factors
10
20
20
20
•lAmortJzationi
0.0672157
0.0943929
0.1174596
3.2.5 Re-Basing Bureau of Labor Statistics Cost Indices
The Water Model and WAV Cost Model uses BLS cost index information based to 1967.
In 1986, the BLS conducted a comprehensive overhaul of the industrial price methodology resulting
in a re-basing of all index information to a 1982=100 base year. This requires a re-basing of BLS
index information to 1967 prior to use in the models for the development of cost estimates. Table
3-11 provides the re-base factors. A sample re-base calculation is presented below.
Sample Re-base Calculation:
Machinery
1982 Base Factor / Re-base Factor = 1967 Base Factor
147.8/0.32895016 = 449.3
3-11
-------
Table 3-11
Bureau of Labor Statistics Re-base Information
^^ft^iif'SIS
Machinery
Concrete
Steel
Pipes & Valves
Electrical
PPI Finish Goods Index
• 1R*fereiitee< :
BLS114
BLS 132
BLS 1017
BLS 1149
BLS 117
BLS 3000
^1982=5l00|;;
-'"'"Number ','
147.8
148.8
113.3
162.2
120.8
130.6
•'^''•IS&ba&e^'
.Vfei^ofe'!'
0.32895016
0.32261652
0.28608856
0.30909034
0.43185069
0.35633299
' i9|7«i6b;;':
>: Number^ ,-'
- 449.3
461.2
396.0
524.8
279.7
366.5
*jBate;rl
9/98
9/98
9/98
9/98
9/98
9/98
> Provided by the BLS
3.2.6 Flows Used in the Development of Costs
Flow categories were developed to provide adequate characterization of costs across each of
the flow regions presented in Section 3.2.1. A minimum of four data points were generated for each
of the flow regions, with the exception of the transition regions, where c;ost estimates are based upon
a linear regressions between the last data point of the previous region £ind the first data point of the
following region. Table 3-12 presents the design and average flows, and cost models used in this
process.
3-12
-------
Table 3-12
Flows Used in the Cost Estimation Process
Design Flow (mgd)
0.010
0.024
0.087
0.10
0.27
0.45
0.65
0.83
... » , 1.0;.' . -:
1.8
4.8
10
11
18
26
51
210
430
Average Flow (mgd)
0.0031
0.0056
0.024
0.031
0.086
0.14
0.23
0.30
0.36
0.7
2.1
4.5
5
8.8
13
27
120
270
Cost Model
vss
vss
vss
; -: 'vss : ,-•;.
Water
Water
Water
Water
Water, ;
WAV Cost
WAV Cost
WAV Cost .
WAV Cost
WAV Cost
WAV Cost
WAV Cost
WAV Cost
WAV Cost
Shaded rows represent data used in the estimation of costs with the transition regions.
3-13
-------
3.3 ADDITIONAL CAPITAL COSTS
The cost models discussed in the previous sections are good tools for estimating capital and
O&M costs associated with various drinking water treatment technologies. There are additional
capital costs, however, which the models do not account for and may be a very real expense for
public water utilities. The need for additional capital costs can be affected by a number of factors,
including: contaminants present, quality of the source water, land availability, retrofit of existing
plants, permitting requirements, piloting issues, waste disposal issues, building or housing needs,
and redundancy. Tables with additional capital cost estimates for each technology discussed in this
document are presented in Appendix E.
Contaminants
Arsenic is typically present in drinking water in one of two oxidation states, As(III) or As(V).
As(V) is more effectively removed by each of the removal technologies discussed in this document.
However, As(III) can be easily oxidized to As(V) using chlorination, potassium permanganate, or
other methods. Groundwaters typically contain As(III), while As(V) is more commonly found in
surface waters.
The presence of additional contaminants, for example, inorganics (sulfate, aluminum,
manganese), pathogenic contaminants (Giardia, Cryptosporidiwri), or organic contaminants
(trihalomethanes, haloacetic acids), can raise additional treatment concerns and result in decreased
process performance. Changes in coagulant dosage or type, sedimentation time, or membrane
efficiency are just a few of the concerns that may arise. Presence of pathogens can result in a need
for disinfection of finished water.
For the purpose of developing cost estimates, it is assumed that the source water is from a
clean, consistent, single source. Facilities combining plant influent from multiple sources can affect
source water quality, which may in turn affect the removal efficiency of the treatment technologies
discussed. Ion exchange, for example, may not be an effective removal technology for source waters
3-14
-------
with high influent sulfate levels.
Costs may vary for ground and surface water systems as well. Treatment technologies
susceptible to fouling by suspended solids (i.e., surface water systems) may require additional
pretreatment (pre-filtration). Costs for these items are included in the additional capital cost tables
presented in Appendix E.
Land
Land requirements were calculated based upon TDP recommendations (EPA, 1997) and
engineering judgement. Appendix E presents two scenarios for land costs. The low cost scenario
assumes land costs to be $1,000 per acre for small systems (i.e., less than 1 mgd) and $10,000 per
acre for large systems. All land costs are $100,000 per acre for the high cost scenario (SAIC, 1998).
Retrofitting
All costs presented in this document are for new construction, with the exception of the
enhanced coagulation and enhanced lime softening processes. All processes contained in the cost
models include pipes and valves, electrical and instrumentation, and other costs associated with
retrofitting. It was assumed that the costs included are sufficient for the retrofit of existing
coagulation/filtration and softening plants. As a result, costs for retrofitting are excluded from
Appendix E.
Permitting
Permitting costs follow the recommendations of the TDP as presented in the Technology
Design Conference Information Package (EPA, 1997). A technology-specific summary of low and
high cost permitting scenarios is presented in Table 3-13. The number of permits required can vary
by location, depending upon State and Local regulations, as well as technology. Some technologies
may require permitting for storage tanks used for process chemicals, while others may necessitate
NPDES permits if the disposal option for process residuals happens to be discharge to a nearby
surface water.
3-15
-------
ja
2
«
H
O)
JS
ft
e
CD
M
S
S
k.
o>
ON
g
U.
Z
u.
X
<
j
U-
^
a.
Q
3_
B
u
S,
Permit Ty
u
'So
0
o
^
g
H
<
-------
Eiloting
Piloting costs are neglected in this document and are not included in Appendix E.
The characteristics of arsenic-containing waste streams is presented in Chapter 4.
Appropriate handling and disposal methods are discussed for residuals generated by each treatment
process for which capital and O&M cost estimates are provided. Cost equations for disposal by each
of these methods are presented in Small Water System Byproducts Treatment and Disposal Cost
Document (DPRA, 1993 a) and Water System Byproducts Treatment and Disposal Cost Document
(DPRA, 1993b).
Storage/Bjulding
All of the cost models used in preparing the technologies and costs document include costs
for housing of equipment. It is assumed that the costs included in the model output is sufficient.
As a result, additional building costs are not included. It is also assumed for all scenarios that source
water production is consistent, and storage for source water is not provided.
Redundancy
The cost models include standby pumps for some of the unit processes used in generating
the cost estimates presented in this document, e.g., raw and finished water pumping. Further, it is
good design practice to include additional filtration structures and sedimentation basins to allow
continued operation during maintenance of one or more of the structures. Backup pumps are not
included for chemical feed systems. As a result, there may be some additional capital costs
associated with redundancy for these items. Recommended Standards for Water Works (Great Lakes
Upper Mississippi River Board of State Public Health and Environmental Managers, 1 997), often
referred to as the Ten State Standards, presents a comprehensive discussion of redundancy and
recommended redundant items. The Ten State Standards were used for presenting costs for
redundant items in Appendix E.
3-17
-------
3.4 COSTS FOR MULTIPLE REMOVAL PERCENTAGES
Capital and O&M cost estimates are presented for the maximum achievable removal in this
document. Table 3-14 presents a removal technology matrix which identifies maximum removal
percentages for the technologies for which costs have been estimated. Costs for facilities requiring
less than the maximum removal to meet the arsenic MCL target, can be estimated using the blending
approach discussed in Section 3.4.2.
3.4.1 Removal and Accessory Costs
Costs for each of the removal technologies presented in this document can be separated into
two categories: removal and accessory. Accessory costs include raw and finished water pumping,
and clearwell storage. Removal costs include any process item directly associated with the removal
of a particular contaminant, e.g., the ion exchange bed in ion exchange processes.
Accessory costs are independent of the desired removal percentage. For example, a one mgd
treatment plant must still pump one million gallons of raw water into the plant, pump one million
gallons of finished water, and have adequate storage (10% of daily production). Conversely,
removal costs are dependent upon the desired removal. If contaminant levels are such that the plant
need only remove 30 percent of the contaminant to reach the treatment goal, then the treatment
process can be scaled to treat a portion of the flow. The treated flow is then blended with the
untreated portion prior to distribution. Section 3.4.2 discusses the blending approach used in the
development of cost estimates.
Cost estimates presented in this document do not include accessory capital and O&M.
Cost curves and equations for accessory costs (i.e., raw and finished water pumping, and clearwell
storage) are presented in Appendix D.
3-18
-------
Table 3-14
Treatment Technology Maximum Achievable Removal Percentages
Treatment Technology
Coagulation/Filtration
Enhanced Coagulation/Filtration1
Direct Filtration
Coagulation Assisted Microfiltration
Lime Softening
Enhanced Lime Softening1
Ion Exchange
Activated Alumina
Reverse Osmosis
Greensand Filtration3
POE Activated Alumina
POU Ion Exchange
Maximum Percent
Removal
95
95
90
90
80
80
95
90
>95
50
70
70
1 - Enhanced processes assume the existing plant can achieve 50% removal without modification.
Process enhancements result in the balance to achieve the maximum removal. For example, an
existing coagulation/filtration facility can achieve 50% removal. Process enhancements result in
an additional 45% removal, for a total removal of 95%.
3.4.2 Use of Blending in Cost Estimates
Capital and O&M costs were estimated vising the VSS, Water, and WAV Cost models. If raw
water contaminant levels are sufficiently low, a utility may not need to achieve maximum removal
to achieve a treatment goal. For example, assume a facility is considering installation of a
coagulation filtration which can achieve 95% removal. If the raw water arsenic concentration is 20
/^g/L and the treatment objective is a finished water concentration of 10 //g/L, the utility need only
remove 50% of the arsenic in the raw water, hi this scenario, the facility could treat a portion of the
raw water and blend with untreated water and still achieve its treatment objective. The portion of
the total process flow to be treated can be calculated using the following equation:
3-19
-------
treated Vtotal
Where: Quoted = Treated portion of the total process flow, mgd
Qwtai = Total daily process flow, mgd
Qnax = Maximum achievable removal efficiency, %
= Desired removal efficiency, %
If 1 is substituted for the total daily flow (Q^^ in the above equation, the treated portion of
the flow (Qtreated) is expressed as a fraction of the total flow. Multiplying that fraction by the total
plant flow will result in design and average operating flows that can be used to estimate capital and
O&M costs for the treated portion of the flow, using the graphs and equations presented in this
Section.
3.5 PRE-OXIDATION PROCESSES
Inorganic arsenic occurs in two primary valence-states, arsenite (As III) and arsenate (As V).
Surface waters more typically contain AsOO, while As(III) is the dominant species found in ground
waters. Each of the treatment technologies presented in this document remove As(V) more readily
than As(III). As a result, pre-oxidation may be necessary depending upon source water conditions.
Potassium permanganate addition and chlorination are two oxidation technologies that have
been evaluated and deemed effective for the conversion of arsenite to arsenate. Chlorination may
cause disinfection by-product (DBF) formation in source waters with high TOC concentrations.
Further, chlorination may cause fouling in some membrane processes. Source water characteristics
should be thoroughly evaluated when considering pre-oxidation technologies. Additional oxidation
technologies, such as ozonation and hydrogen peroxide, may be effective, but need further
evaluation.
3.5.1 Potassium Permanganate
Potassium permanganate can be used as a pre-oxidation technology for conversion of As(III)
to As(V). Potassium permanganate is more expensive than chlorination; $2700 per ton compared
with $350 to $400 per ton. However, unlike chlorination, potassium p
-------
form measurable DBFs and does not foul membranes. Raw water and downstream process
considerations should be made when selecting a pre-oxidation technology. For this document
potassium permanganate costs were calculated for dosages of 1.5,3.0 and 5.0 mg/L. The Very Small
Systems Best Available Technology Cost Document (Malcolm Pirnie, 1993) was used for calculating
costs for the flows below 1 mgd. For flows greater than 10 mgd the W/W Cost Model was used to
estimate the capital and O&M costs. Linear regressions were used to estimate costs in the transition
regions between the two models, i.e., 1 to 10 mgd. The following are some highlights of the system
design used at the time of cost estimation:
• For very small systems, the potassium permanganate feed system is equipped with a
metering pump, solution tank with mixer, pipes and valves, and instrumentation and •
controls. The system utilizes a 3% potassium permanganate solution.
• The VSS document makes provisions for building (42.7%), fencing (49.4%), and road
(33.8%) costs associated with potassium permanganate addition. However, costs
presented in this document do not include these items. It is assumed permanganate
addition will be installed as part of a larger treatment process and that building, fence and
road costs for the treatment facility are adequate to accommodate permanganate additon.
• O&M costs for very small systems were calculated using equations in the very small
systems cost document. Labor requirements were assumed to be 1 hour per week.
• For small system potassium permanganate addition, a dry chemical feed system capable
of 1,000 pounds per day was used.
Figures 3-1 through 3-6 present capital and O&M cost curves and equations for potassium
permanganate addition.
3-21
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3.5.2 Chlorination
As previously stated, chlorination can cause DBF formation in source waters with high TOC
concentrations. Chlorination has also been shown to cause fouling in some membrane processes.
As a result, source water characteristics and downstream process needs should be thoroughly
evaluated when considering chlorination as an oxidation technology. Capital and O&M costs were
developed for cylinder and tank feed chlorination systems at dosages of 1.5, 3.0, and 5.0 mg/L.
Similar to potassium permanganate systems, the VSS Model was used for calculating costs for the
flows below 1 mgd, and the WAV Cost Model was used to estimate the: capital and O&M costs for
flows greater than 10 mgd. Linear regressions were used to estimate costs in the transition regions
between the models, i.e., 1 to 10 mgd. The following are some highlights of the system design used
at the time of cost estimation:
• For very small systems, chlorination is accomplished with a hypochlorite feed system
capable of providing dosages to 10 mg/L as chlorine. The system is equipped with a 150
gallon storage tank and utilizes a 15% sodium hypochlorite: feed stock.
• The VSS Model makes provisions for building (52.2%), fencing (60.5%), and road
(41.4%) costs associated with chlorine addition.
• Capital costs were calculated for both with and without housing costs added.
• Labor requirements for O&M costs were assumed to be 1 hours per week.
• For small systems, cylinder feed chlorination system capital and O&M costs were
estimated.
It should be noted that some systems currently using chlorine for disinfection may be able
to modify existing chlorine feed systems to utilize chlorine as a preoxidant with significant capital
cost savings. Capital and O&M cost curves and equations are presented in Figures 3-7 through 3-12.
3-28
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3.6 PRECIPITATIVE PROCESSES
3.6.1 Coagulation/Filtration
Coagulation/filtration (C/F) is a treatment process that alters the physical or chemical
properties of colloidal or suspended solids, enhancing agglomeration, and allowing these solids to
settle out of solution by gravity or be removed by filtration. The C/F removal mechanism is
discussed in greater detail in Chapter 2. A typical C/F process includes coagulant addition, which
may be followed by polymer addition to aid agglomeration, flocculation, sedimentation, and
filtration.
C/F is widely used as a treatment for removing suspended solids from surface water supplies.
Most ground waters are low in turbidity and do not require this type of treatment. Source waters
containing high As(III) concentrations may opt for oxidation as a pre-treatment for C/F. Pre-
oxidation options are presented in Section 3.5.
Capital costs for very small systems were developed using the VSS model. The design
parameter most affecting capital cost is the filtration rate. It affects the size of the filter structure and
volume of filter media, the most cost intensive process in a C/F plant. The VSS model also makes
provisions for building (14.9-28.1%), fencing (2.1-7.5%) and road (1.2-4.7%) costs associated with
each of the technologies presented. The following design criteria were used to develop capital cost
estimates for systems with a design flow of less than 0.10 mgd:
• Coagulant dosage, ferric chloride, 25 mg/L;
• Polymer dosage, 0.4 mg/L; and
• Filtration rate, 2.5 gpm/ft2.
O&M costs are most affected by chemical costs associated with coagulant and polymer
dosages. As a result, the very small systems O&M cost estimates were escalated using the BLS
Chemical and Allied Products Index. Labor requirements were estimated at 8 hours per week.
3-35
-------
Small Systems (Less than 1 mgd)
The Water Model was used to estimate capital and O&M costs for small C/F treatment
plants. The following design criteria were used in developing capital and O&M cost estimates:
• Package plant for all small systems, filtration rate 5 gpm/fl:2; .
• Ferric chloride dose, 25 mg/L;
• Polymer dose, 2 mg/L; and
• Lime dose, 25 mg/L for pH adjustment.
Large Systems (Greater Than 1 mgd)
The WAV Cost model was used to develop capital and O&M cost estimates for large C/F
plants. The following design criteria were used to estimate capital and O&M costs:
Ferric chloride dose, 25 mg/L;
Polymer dose, 2 mg/L;
Lime dose, 25 mg/L for pH adjustment;
Rapid mix, 1 minute;
Flocculation, 20 minutes;
Sedimentation, 1000 gpd/ft2 using rectangular tanks; and
Dual media gravity filters, 5 gpm/ft2.
Figures 3-13 and 3-14 present capital and O&M cost curves and equations for removal of
arsenic by C/F.
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3-38
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3.6.2 Enhanced Coagulation
Enhanced coagulation involves modifications to the typical C/F process such as increasing
the coagulant dosage, reducing the pH, or both. The process is nearly identical to that of
conventional C/F with those two exceptions. Source waters with high influent As(III) concentrations
may require pre-oxidation for conversion of arsenite to arsenate (see Section 3.5).
For the purpose of estimating costs, it was assumed that a typical C/F treatment plant could
remove 50 percent of the influent arsenic prior to modification, i.e., enhancement. It was also
assumed that the only added O&M burden would result from power and materials costs, no
additional labor was assumed to be required. Costs presented are for the enhancement only, and
are in addition to any current annual debt incurred by the utility.
The VSS Model makes no appropriations for estimating enhanced coagulation capital and
O&M costs. As a result, the Water Model was used to estimate capital and O&M costs for all
enhanced coagulation treatment plants with a capacity of less than 1 mgd. The following design
criteria were used in developing capital and O&M cost estimates:
• Additional ferric chloride dose, 10 mg/L;
• Additional feed system for increased ferric chloride dose;
• Additional lime dose, 10 mg/L for pH adjustment; and
• Additional feed system for increased lime dose.
Large Systems (Greater Than 1 mgd)
The W/W Cost model was used to estimate capital and O&M costs for large enhanced
coagulation plants. The following design criteria were used to estimate costs:
• Additional ferric chloride dose, 10 mg/L;
• Additional feed system for increased ferric chloride dose;
• Additional lime dose, 10 mg/L for pH adjustment; and
• Additional feed system for increased lime dose.
Figures 3-15 and 3-16 present capital and O&M cost curves and equations for enhanced
coagulation.
3-39
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3-41
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3.6.3 Direct Filtration
Direct filtration is a modified C/F treatment process utilized for source waters with low
influent suspended solids concentrations. Because of the low solids content, settling is not required
and coagulation is followed immediately by filtration. Direct filtration includes all of the typical C/F
process elements with the exception of flocculation and sedimentation. Source waters with high
influent As(III) concentrations may require pre-oxidation for conversion of arsenite to arsenate. Pre-
oxidation technologies are discussed in Section 3.5.
Capital costs for very small systems were developed using the VSS Model. The design
parameter most affecting capital cost is the filtration rate. This paramet<;r affects the size of the filter
structure and volume of filter media, the most cost intensive processes within a direct filtration plant.
The VSS Model also makes provisions for building (23.5-1 9.5%), fencing (3.2 - 7.5 %) and road (2.0
- 5.0 %) costs associated with each of the teclmologies presented. The following design criteria were
used to develop capital cost estimates for very small direct filtration systems:
• Coagulant dosage, alum or ferric chloride, 20 mg/L;
• Polymer dosage, 1 .0 mg/L; and
• Filtration rate, 5.0 gpm/ft2.
O&M costs are most affected by chemical costs associated with coagulant and polymer
dosages. As a result, the very small systems O&M cost estimates were escalated using the BLS
Chemical and Allied Products Index. Note that labor requirements were estimated at 8 hours per
week.
Stna|l Systpins (Less than 1 itigd)
The Water Model was used to estimate capital and O&M costs for small direct filtration
plants. The following design criteria were used in developing capital and O&M cost estimates:
• Ferric chloride dose, 10 mg/L;
» Polymer dose, 1 mg/L;
• Lime dose, 1 0 mg/L for pH adjustment.
3-42
-------
T.arge Systems (Greater Than 1 mgd)
The WAV Cost model was used to develop capital and O&M cost estimates for large direct
filtration plants. The following design criteria were used to estimate costs:
Ferric chloride dose, 10 mg/L;
Polymer dose, 1 mg/L;
Lime dose, 10 mg/L for pH adjustment;
Rapid mix, 1 minute;
Flocculation, 20 minutes; and
Dual media gravity filters, 5 gpm/ft2.
Figures 3-17 and 3-18 present capital and O&M cost curves and equations for removal of
arsenic by direct filtration.
3-43
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3-45
-------
3.6.4 Coagulation Assisted Microfiltration
Coagulation assisted microfiltration is another modified C/F process wherein microfiltration
is used in place of a conventional gravity filter. The process includes all coagulation, flocculation,
sedimentation, and microfiltration. Coagulation assisted microfiltration is capable of removing
smaller particle floe which results in decreased coagulant dosage and increased plant capacity.
Source waters with high influent As(III) concentrations may require pre-oxidation for conversion
of arsenite to arsenate. Pre-oxidation technologies are discussed in Section 3.5.
Capital and O&M costs for very small systems for the coagulation portion of this process
were developed using the VSS Model. The C/F design parameters giv«ai in Section 3.6. 1 were used
here as well, however no polymer addition was assumed due to the tendency of some membranes
to foul when used with polymer. Very small system costs also include a 20 mg/L sodium hydroxide
dose to adjust process pH. Microfilter specifications and cost estimates were developed based upon
vendor quotes and case studies. These costs were then added to the C/F cost estimates.
SmaUJSyjsJsjmsj(LessJhanJLmgd)
The Water Model was combined with vendor data and case studies to estimate capital and
O&M costs for coagulation assisted microfiltration treatment plants. The following design criteria
were used in developing capital and O&M cost estimates for the coagulation portion of this process:
• Package plant for all small systems, filtration rate 5 gpm/ft2;
• Ferric chloride dose, 25 mg/L;
• Sodium hydroxide dose, 20 mg/L; and
• Standard microfilter specifications, provided by vendors.
3-46
-------
Large Systems (Greater Than 1 mgd)
The WAV Cost model was used to develop capital and O&M cost estimates for large
coagulation assisted microfiltration plants. The following design criteria were used to estimate
capital and O&M costs for the coagulation portion of this process:
• Ferric chloride dose, 25 mg/L;
• Rapid mix, 1 minute;
• Flocculation, 20 minutes;
• Sedimentation, 1000 gpd/ft2 in rectangular basins; and
• Standard microfilter specifications, provided by vendors.
Figures 3-19 and 3-20 present capital and O&M cost curves and equations for removal of
arsenic by coagulation assisted microfiltration.
3-47
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3-49
-------
3.6.5 Lime Softening
Lime softening (LS) has been widely used for reducing hardness in large water treatment
systems. The LS removal mechanism is discussed in Chapter 2. As(III) or As(V) removal by LS
is largely pH dependent, and pre-oxidation of arsenite to arsenate will significantly improve arsenic
removal efficiencies. Pre-oxidation technologies are discussed in Section 3.5.
Considerable amounts of sludge are produced by the LS process. Large systems may find
it economically feasible to install recalculation equipment to recover and reuse the process sludge
and reduce disposal costs.
SmalLS^Jtem^LessJhanJLjngd)
The VSS Model provides no estimation methods for LS treatment. Therefore, the Water
Model was used to estimate capital and O&M costs for all LS treatment plants with less than 1 mgd
capacity. The following design criteria were used in the development of capital and O&M cost
estimates:
• Package plant for all small systems;
• Lime dose, 250 mg/L; and
• Carbon dioxide (liquid), 35 mg/L for recarbonation.
The WAV Cost model was used to develop capital and O&M cost estimates for large LS
plants. The following design criteria were used to estimate capital and O&M costs:
Lime dose, 250 mg/L;
Carbon dioxide (liquid), 35 mg/L for recarbonation;
Rapid mix, 1 minute;
Flocculation, 20 minutes;
Sedimentation, 1500 gpd/ft2 using circular tanks; and
Dual media gravity filters, 5 gpm/ft2.
Figures 3-21 and 3-22 present capital and O&M cost curves and equations for lime softening.
3-50
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3-52
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3.6.6 Enhanced Lime Softening
Enhanced LS involves modifications to the typical LS treatment process in the form of
increased lime dosage and possibly increased soda ash dosage. This may result in the need for pH
adjustment of treated water via recarbonation. Source waters with high influent As(III)
concentrations may require pre-oxidation for conversion of arsenite to arsenate. Pre-oxidation
technologies are discussed in Section 3.5.
For the purpose of estimating costs, it was assumed that an existing LS plant could achieve
50 percent removal of arsenic from the source water prior to modification, i.e., enhancement. It was
also assumed that the only added O&M burden would result from power and materials costs, no
additional labor was assumed to be required. Costs presented are associated with the
enhancement only, and are in addition to current annual debt incurred by the utility.
SjnialLS^siejns^LessJdbLanLLmgd)
The Water Model was used to estimate capital and O&M costs for small enhanced LS
treatment plants. The following design criteria were used in the development of capital and O&M
cost estimates:
• Additional lime dose, 50 mg/L;
• Chemical feed system for increased lime dose;
• Additional carbon dioxide (liquid), 35 mg/L for recarbonation; and
• Chemical feed system for increased carbon dioxide dose.
Large Systems (Greater Than 1 mgd)
The W/W Cost model was used to develop capital and O&M cost estimates for large
enhanced LS plants. The following design criteria were used to estimate capital and O&M costs:
• Additional lime dose, 50 mg/L;
• Chemical feed system for increased lime dose;
• Additional carbon dioxide (liquid), 35 mg/L for recarbonation; and
• Chemical feed system for increased carbon dioxide dose.
Figures 3-23 and 3-24 present cost curves and equations for removal of arsenic by enhanced
lime softening.
3-53
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3-55
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3.7 ADSORPTION PROCESSES
3.7.1 Activated Alumina
Activated Alumina (AA) is a physical/chemical process by which ions in the feed water are
sorbed to the oxidized AA surface. Feed water is continuously pass
-------
Table 3-16
Influent pH vs. Regeneration for Activated Alumina
Influent pH
7.5
7
6.5
6
Bed Volumes Treated to
10 ng/L limit
9,000
20,000
57,500
110,500
As shown by these results, the number of bed volumes is highly dependent on pH. Due to
insufficient information available in the literature, the cost equations given in this document do not
take influent pH into account. However, when estimating the cost for a particular system, the
influent arsenic concentration and the influent pH should be taken into account when possible.
The very small systems costs for activated alumina processes were generated by the EPA
based on the information given in Appendix G. These costs were assembled from the VSS Model,
the Water Model, and two draft ORD reports. These costs were treated slightly differently than other
VSS costs due to two primary assumptions made by the EPA:
• Very small systems will not provide pH adjustment, but will treat water at the ambient
pH.
• Very small systems will not regenerate, but will dispose of the AA media when a bed is
exhausted.
Because these assumptions were not made in either the VSS or Water models, the EPA adjusted
the VSS costs to account for them. Capital and O&M costs were calculated for very small systems
assuming no regeneration of the AA medium. These curves and equations are provided in Figures
3-25 and 3-26.
3-57
-------
Small Systems (Less than 1 mgd)
For systems less than 0.66 mgd, the same procedure outlined for very small systems was
employed by the EPA to generate costs based on the same information and assumptions included
in Appendix G.
For systems between 0.66 and 1 mgd, the Water Model was used to estimate capital and
O&M costs for small treatment plants. Plants larger than 0.66 mgd ware assumed to lower influent
pH for optimal operation. O&M costs were determined in the same manner as described above for
the very small systems. The design criteria used in development of costs are:
• EBCT of 1 5 minutes;
• Sulfuric acid feed, 70 mg/L;
• Caustic feed, 28.5 mg/L;
• Regenerant dose of 0.3 Ib NaOH/kgal
The WAV Cost model was used to determine the activated alumina capital and O&M costs
for large systems. The following design criteria were assumed for determination of AA costs:
Sulfuric acid feed, 70 mg/L;
Caustic feed, 28.5 mg/L;
Operation at pH 5.5;
Low influent suspended solids;
10-foot deep beds;
100 psi working pressure;
80 percent bed expansion during backwash;
Regeneration storage facilities sized for 30-day requirement;
NaOH in solid phase for plants less than 10 mgd;
NaOH in 50% solution for plants greater than 10 mgd;
Media replacement of 10% per year.
The WAV Cost model assumes the number of bed volumes to breakthrough is 1600. The
O&M costs for other BVs is calculated in proportion to the assumed BV of 1 600. The WAV Cost
model was also used to calculate capital and O&M costs associated with acid and base feed. Figures
3-58
-------
3-27 through 3-36 present capital and O&M cost curves and equations for AA. The O&M costs
were calculated for BVs of 500,2000,3000, 5000, 7000,10000, 16500,25000, and 50000.
3-59
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3-71
-------
3.8 ION EXCHANGE PROCESSES
3.8.1 Anion Exchange
Ion exchange (IX) is a physical/chemical treatment process in which an ion on the solid phase
(DC resin) is exchanged for an ion in the feed water, thereby removing contaminants from the feed
water. The IX removal mechanism is discussed in detail in Chapter 2. Ion exchange resin can be
fouled by suspended and dissolved contaminants in the feed water. If the feed water contains
suspended solids the IX process will need to be preceded by a pretreatment process, typically multi-
media filtration. Also, source waters high in As(III) concentration may require pre-oxidation for
conversion of arsenite to arsenate. Pre-oxidation is discussed in Section 3.5. Neither pre-oxidation
nor pre-filtration have been considered as part of the costs develop**! in this section.
Sulfate concentrations hi the influent water significantly affect the capacity of the IX resin
with respect to the removal of arsenic. Clifford (1993) estimated bed volumes for 10 percent and
50 percent breakthrough of influent arsenic as a function of influent sulfate concentration. Figure
3-37 shows the bed volume and sulfate relationship estimated by Clifford (1993). Figures 3-38
through 3-41 were developed using the relationship shown in Figure 3-37. Using these figures, the
regeneration frequency for an IX column can be estimated if the inflvient arsenic, sulfate, and target
effluent arsenic concentrations are known. Straight line fits of the data points derived from Figure
3-37 are also shown on figures 3-38 through 3-41. Once the BV is known, the corresponding
equation may be used to estimate the O&M cost. Capital costs will not be affected by changes in
bed volume to regeneration.
3-72
-------
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EPA developed a modified capital cost estimation method based upon the VSS Model and
the Water Model (Kempic, 1994d). This method was used to develop the very small systems capital
costs presented in this document (see Appendix F). The ion exchange plant was assumed to be
comprised of the following components:
Pressure ion-exchange unit;
Brine dilution pump;
Brine pump;
Salt storage tank;
Pipes and valves;
Instrumentation and controls.
The referenced EPA cost approach assigns percentages to each of the construction cost
components (i.e., sitework, equipment, concrete, steel, labor and installation, pipes and valves,
electrical, and housing) and escalates each component using the appropriate BLS cost index. The
following design parameters were used in the development of the very small systems capital costs:
• EBCT of 2.5 minutes;
• Regenerant dose of 1 5 lb/ft3 of resin;
• Regenerant frequency of once per day.
O&M costs for the ion exchange process were based on the O&M cost equation given in the
VSS Model. The use of these equations was confirmed in the revised EPA cost method presented
above (Kempic, 1994d). The document assumed 576 bed volumes were treated per day. Based on
this number of bed volumes, regeneration frequency was calculated for each assumed number of bed
volumes to breakthrough. The regeneration frequency was used as an input to the O&M cost
equation. The assumed labor requirement was 1 0 hours per week.
Small Sir
The Water Model was used to estimate capital and O&M costs for small treatment plants.
For the O&M costs, the Water Model assumes daily regeneration of the DC bed. However, the
number of bed volumes treated per day is not given for this model. Therefore, a bed volume to
3-78
-------
breakthrough of 160 was assumed since this was the same number assumed in the WAV Cost
program for large systems. The O&M costs for each bed volume to breakthrough were then
calculated in proportion to the assumed regeneration frequency of 160 bed volumes.
The resin cost assumed in the Water Model was $210/ft3. A survey of ion exchange facilities
and suppliers found this unit cost to be high. Based on this survey, $125/ft3 was determined to be
an appropriate unit cost. Therefore, ion exchange costs were adjusted to reflect a change of $210/ft3
to 5125/ft3. Additionally, input parameters were set in the Water Model to adjust for an EBCT of
2.5 minutes. A sensitivity analysis was performed to determine the effect of reducing EBCT to 1.5
minutes from 2.5 minutes.
Large Systems (Greater than 1 mgd)
The WAV Cost model was used to determine the ion exchange capital and O&M costs for
large systems. The following design criteria were assumed for determination of IX costs:
Resin cost of $125/ft3;
EBCT of 2.5 minutes;
Six-foot deep DC beds;
100 psi working pressure;
Nitrate =100 mg/L;
Sulfate = 80 mg/L;
Other anions =120 mg/L;
Nitrate capacity = 7 kilograins/ft3 resin;
Regenerant requirement = 15 Ib NaCl/ft3;
Regeneration time = 54 minutes;
Backwashing time =10 minutes;
Rinsing time = 24 minutes;
25% resin replacement per year.
In the WAV Cost model, the number of bed volumes to breakthrough is assumed to be 160.
The O&M costs for other BVs is calculated in proportion to the assumed regeneration frequency of
160 BVs. A sensitivity analysis was performed to determine the effect of reducing EBCT to 1.5
minutes from 2.5 minutes. This reduction would have resulted in about a 15 percent reduction in
costs.
3-79
-------
Figures 3-42 through 3-47 present capital and O&M cost curves and equations for ion
exchange. All costs include redundant exchange beds, i.e., beds for use during maintenance of
another that keep the system on-line. Table 3-17 presents the number of DC beds included in the cost
estimates provided.
Table 3-17
Number of EX Beds Included in Cost Estimates1
Plant Capacity (mgd)
<1.2
1.2-3.9
3.9 - 6.5
6.5 - 13
>13
Number of IX Beds
2
3
5
10
10
1- Culp/Wesner/Culp (1979,1984)
3-80
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3-86
-------
3.9 SEPARATION PROCESSES
3.9.1 Microfiltration
Microfiltration is a low-pressure membrane process which has only a marginal ability to
remove arsenic due to its relatively large pore size in comparison to other membrane processes. MF
removes contaminants from a feed stream primarily through sieving. Typically, MF does not require
pretreatment beyond approximately 500-um prefiltration. Because MF is not an effective stand-
alone technology for removal of arsenic, capital and O&M cost estimates for MF are not provided
in this chapter. Appendix D, however, does contain estimated costs for MF treatment.
The Water and WAV Cost models used in determination of other process costs in this
document do not include MF. Although the VSS Model did include MF, the equations were only
used as a final check for reasonableness of costs for systems with a capacity less than 0.10 mgd.
Since MF is a relatively new technology, very little cost information is contained in the literature.
However, the MF cost information which currently exists provided the basis for the estimates
contained in Appendix D. This information consisted of cost surveys from existing plants and
vendor quotes. Vendor quotes were obtained from Memtec, a manufacturer of package MF systems.
Cost survey information was obtained from several sources. The primary source was a survey of 21
MF plants conducted as part of a 1996 AWWARF study.
The largest MF plant which provided cost information had a capacity of 20 mgd. For this
reason, the economies-of-scale which exist beyond a capacity of 20 mgd could not be accurately
estimated. Therefore, no economies-of-scale were assumed beyond the boundary condition of 20
mgd for either capital or O&M costs.
3.9.2 Ultraffltration
Ultrafiltration is a low-pressure membrane process which removes contaminants from a feed
stream primarily through sieving. Typically, UF does not require pretreatment beyond
approximately 200-um prefiltration. UF has the benefit of being lower in both capital and O&M
costs than high-pressure membrane processes.
3-87
-------
Although the WAV Cost and VSS models included UF, only the WAV Cost model was used
for UF cost calculations. The very small systems cost equations were used only as a check against
final UF costs. The WAV Cost estimation was valid to a capacity of 1 mgd. Since the time when
the WAV Cost model was assembled, however, UF membrane module costs have decreased by
approximately 30 percent. For this reason, the membrane module piortion of the capital costs was
reduced by 30 percent to account for this. Also, the membrane replacement portion of the O&M
costs was reduced by 30 percent to account for this as well. Actual plant cost information was also
used for UF cost estimation. Since UF is a new technology, however, very little cost information
is contained in the literature. The UF cost information which cuiTently exists was used in the
calculation of the UF capital and O&M costs. This information consisted of cost surveys from
existing plants. The primary source of cost information was a survey of seven UF plants conducted
as part of an AWWARF study in 1996.
The largest UF plant which provided cost information had a capacity of 28 mgd, and the
economies-of-scale which exist beyond a capacity of 28 mgd could not be accurately estimated.
Therefore, no economies-of-scale were assumed beyond the boundary condition of 28 mgd for either
capital or O&M costs.
Figures 3-48 and 3-49 present capital and O&M cost curves and equations for UF.
3-88
-------
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-------
3.9.3 Nanofiltration
Nanofiltration is a high-pressure membrane process capable of significant arsenic removal.
NF removes contaminants from a feed stream primarily through a combination of diffusion and
sieving mechanisms. Typically, NF requires pretreatment to remove suspended solids and other
foulants from the feed stream. NF has greater arsenic removal capabilities than low-pressure
membrane processes, however, capital and O&M costs for NF are usually greater than equivalent
costs for low-pressure processes.
The WAV Cost and Water models used in the determination of other process costs in this
document did not include NF. Although the VSS Model did include NF, the cost equations were
used only as a check against final NF costs. Since NF is a relatively new technology, very little cost
information is contained in the literature. However, the NF cost information which currently exists
provided the basis for the NF cost calculations given in this document. This information consisted
of cost surveys from existing plants (Bergman, 1996) and vendor quotes. A summary of how the
cost estimates were derived from this cost survey data is given below for large and small systems.
Synall Systems (Less than 1 mgd)
Capital costs for small systems, i.e. less than 1 mgd, were determined by contacting vendors
to obtain membrane cost estimates for small systems. The cost information that was obtained
consisted of costs for complete treatment systems. This was done in 1992 during the regulatory
negotiations of the DBF Rule as part of the work conducted by the Technologies Working Group
(TWO). The TWO concluded that these costs were representative of small system NF costs. For
this document, the 1992 costs were updated to 1998 dollars using the BCI.
Large Systems ((Greater than 1 mgd)
Capital and O&M cost estimates for large systems were developed from NF cost data
presented in a NF plant survey conducted by Bergman (1996). Costs presented hi this survey were
escalated to 1997 dollars using the 1997 and 1995 ENR Building Cost Indices. Capital and O&M
costs presented here were derived from cost data submitted by existing plants. The Bergman survey,
however, presented capital cost and O&M cost data obtained between 1988 and 1996. It is
3-91
-------
recognized that spiral-wound membrane modules, which include the majority of NF membranes,
have decreased in cost significantly in recent years. Based on vendor information, costs for spiral-
wound membrane modules have been reduced by approximately 50 percent over the past five years.
For this reason, costs for membrane modules presented in the above reference obtained between
1988 and 1995 were reduced by 50 percent. The reduced cost items included new membrane capital
costs and O&M membrane replacement costs.
The costs given in the Bergman survey consist solely of nanofiltration costs from Florida
plants. The source waters treated by these plants are warm, resulting in higher membrane flux values
than potential flux values for lower temperature waters of comparable quality. As a result, the costs
presented in these references may not be representative of costs for all areas of the country. For this
reason, the costs were adjusted to equivalent costs at 20 degrees Celsius. This was accomplished
by assuming a temperature of 25 degrees Celsius for the Florida plants! and adjusting the membrane
capital costs and O&M membrane replacement costs to account for the additional membrane area
that would be required at a lower temperature. The temperature correction equation for permeate
flux [ VJ25=1 .OS01"-250 ] was used for these calculations (Wiesner and Aptel, 1996).
Best-fit curves were generated for capital and O&M costs. Capital costs were separated into
membrane module costs and facility costs. Using the Bergman survey data, the average facility cost
was found to be between two and three times the cost for membrane equipment. However, capital
and operational costs for clearwells and high service pumping will not be required in a retrofit
situation. Subtracting the capital costs for these two components results in a factor of 1.5 to 2.0 for
facility costs when compared to membrane system costs. For this reason, facility costs used in
determining the best-fit equation were calculated by multiplying the membrane cost for each plant
by two (a conservative estimate). It should be noted that the largest plant surveyed was 14 mgd.
Since at the present time very few facilities above this capacity exist, there is no way to accurately
judge the economies-of-scale that may be seen beyond this point. For this reason, it was
conservatively assumed that no economies-of-scale would exist beyond 14 mgd.
Figures 3-50 and 3-51 present capital and O&M cost curves arid equations for NF.
3-92
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3.9.4 Reverse Osmosis
Reverse Osmosis is a high-pressure membrane process which removes dissolved
contaminants from a feed stream primarily through diffusion rather than physical straining. RO
requires a high quality feed stream and often requires substantial pretreatment to remove suspended
solids and other foulants. RO also often requires pH adjustment after the membrane process and
may require the addition of an anti-sealant before the membrane process. For the purpose of this
analysis, costs were not provided for a substantial pre-treatment system, other than the anti-sealant
system. RO has the benefit of greater arsenic removal compared to low-pressure membrane
processes, but is typically associated with higher capital and O&M costs.
Both the VSS Model and the WAV Cost Model included cost estimation for RO. Since the
WAV Cost Model was assembled, however, RO spiral-wound membrane module costs have
decreased by approximately 50 percent. For this reason, the membrane module portion of the capital
costs was reduced by 50 percent. The membrane replacement portion of the O&M costs was also
reduced by 50 percent to account for reductions in membrane costs. The WAV Cost Model for RO
was only valid up to a capacity of 200 mgd. For this reason, no economies-of-scale were assumed
for plants with a capacity larger than the boundary condition of 200 mgd. The model also makes an
assumption that recovery is 80% for systems of 1 to 10 mgd, and 85% for systems larger than 10
mgd. Costs were adjusted to reflect a recovery of 75%.
Figures 3-52 and 3-53 present capital and O&M cost estimates for RO.
3-95
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3-97
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3.10 GREENSAND FILTRATION
Greensand filtration is an oxidation filtration process that has demonstrated effectiveness for
the removal of arsenic. The greensand filtration medium is produced by treating glauconite sand
with KMnO4 until the granular material (sand) is coated with a layer of manganese oxides,
particularly manganese dioxide. Arsenic compounds displace species from the manganese oxide
(presumably OH' and H2O), becoming bound to the greensand surface - in effect an exchange of
ions. The oxidative nature of the manganese surface converts As(III) to As(V), and As(V) is
adsorbed to the surface.
The VSS model was used for estimating greensand filtration capital and O&M costs. This
technology is considered to be a small systems technology and as a result costs were not estimated
bCosts are based upon the following design and operating criteria:
• Potassium permanganate feed, 10 mg/L;
• The filter medium is contained in a ferrosand continuous regeneration filter tank
equipped with an underdrain;
• Filtration rate, 4 gpm/ft2; and
• Backwash is sufficient for 40 percent bed expansion.
Figures 3-54 and 3-55 present cost estimates forjremoval of arsenic by greensand filtration.
3-98
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3.11 COMPARISON OF COSTS
Capital and O&M cost estimates in this chapter were compared with actual cost data
presented in Evaluation of Full-Scale Treatment Technologies at Small Drinking Water Systems
(ICF and ISSI, 1998) and estimates found in Evaluation of Central Treatment Options as Small
System Treatment Technologies (SAIC, 1999). It was found that the cost estimates presented in this
document are generally within the range of values of previous estimates and actual data. Capital cost
estimates were routinely conservative, but followed the general trends seen in actual data, and O&M
estimates typically represented an approximate average of the actual costs. Actual data were not
available for all technologies, and as a result, comparisons are not presented for some of the
technologies discussed in this document
Figures 3-56 through 3-71 provide graphical representations of comparisons between the cost
estimates in this document (labeled "ICI, 1999" on the figures) and actual cost data ("Actual Data"
on the figures). Also, where applicable, cost data are included from: (1) the 1993 T&C document
("EPA, 1993"); and (2) the EPA report, Evaluation of Central Treatment Options as Small System
Treatment Technologies ("SAIC, 1999") prepared by SAIC (1999). The ICI 1999 curves are based
on the cost equations presented earlier in this chapter. For the purpose of comparison, it was
assumed that 100 percent of the process flow is treated.
3.11.1 Capital Cost Comparison
Capital cost estimates were routinely conservative and followed the general trends seen in
actual data. This trend is noted in the comparison of capital cost data for coagulation/filtration. The
majority of actual data points fall below the SAIC 1999 and ICI 1999 curves, indicating a
conservative estimate projected by both SAIC and ICI. Data reported by EPA in 1993 for the
coagulation/filtration technology tends to conform with the ICI 1999 data.
The majority of actual data projected for direct filtration also fall below the SAIC 1999 and
ICI 1999 curves, indicating a conservative estimate. No direct filtration data were developed by
EPA in 1993.
3-101
-------
One actual data point was available for lime softening capital cost comparison. Although
there is a sparsity of actual data, the similarity of the SAIC and ICI curves suggests that the curves
are a realistic representation of actual capital costs.
Three actual data points were available for the activated alumina capital cost comparison.
The three data points fall within a reasonable distance of both the SAIC and ICI curves, but the ICI
curves are more conservative.
Ion exchange capital cost curves developed by SAIC and ICI are similar, especially for larger
systems, and are close to the "Actual Data" and the EPA 1993 data.
Very little actual data exist for ultrafiltration and nanofiltration. Comparisons of ICI 1999
and SAIC 1999 capital costs show similar values, suggesting realistic estimates.
The reverse osmosis comparison of capital costs shows that the; majority of actual data points
fall below the SAIC 1999 and ICI 1999 curves, indicating a conservative estimate projected by both
SAIC and ICI.
In many cases EPA 1993 estimates are greater than the ICI 1999 estimates. This is
because the 1993 estimates included accessory costs, i.e., raw and finished water pumping and
clearwell storage.
3.11.2 O&M Cost Comparison
O&M estimates typically represent an approximation of lie actual costs among the
technologies discussed. The O&M cost curve developed in this document for coagulation/filtration
illustrates that trend. Although there is a good deal of scatter in the actual data points, many of them
cluster around the ICI 1999 curve, especially for the smaller systems. The SAIC 1999 curve falls
below the actual data and tends to be less conservative than the ICI 1999 curve.
The single actual data point for direct filtration O&M cost falls near the ICI 1999 curve.
Both the ICI and SAIC curves show a similar trend and values, with the ICI curve being a bit more
conservative.
The actual data points for O&M costs for lime softening cluster around both the ICI and
SAIC curves, suggesting the validity of these cost estimates, especially for small systems.
3-102
-------
Two actual data points were available for the activated alumina O&M cost comparison. The
O&M cost curves developed by both SAIC and ICI project a conservative estimate when compared
to these two data points.
The actual data points for ion exchange O&M lie close to the ICI 1999 curve, with the SAIC
curve being a bit more conservative, especially for larger systems.
Again, actual data values were not available for nanofiltration and ultrafiltration O&M costs.
The SAIC 1999 estimate was generally more conservative for ultrafiltration and less conservative
for nanofiltration than the ICI 1999 curves. The EPA 1993 data for nanofiltration closely matched
the ICI curve.
The comparison of reverse osmosis O&M cost curves indicates that both curves are close to3
but generally higher than, the actual data points. The ICI 1999 curve is slightly more conservative
than the SAIC 1999 curve.
3-103
-------
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4.0 RESIDUALS HANDLING AND DISPOSAL ALTERNATIVES
4.1 INTRODUCTION
Each of the treatment technologies presented in Chapter 3 will produce residuals, either solid
or liquid streams, containing elevated levels of arsenic. It is the purpose of this chapter to present
the characteristics of arsenic waste generated by each of the treatment technologies and discuss
appropriate handling and disposal options. Costs for residuals handling and disposal are not
presented in this chapter, however, references which contain appropriate cost information are noted.
4.1.1 Factors Affecting Residuals Handling and Disposal Costs
There are a number of factors which can influence residuals handling and disposal costs.
This discussion is concerned with factors affecting capital cost, as well as factors affecting
operations and maintenance (O&M) costs. Capital costs include equipment, construction,
installation, contractor overhead and profit, administrative and legal fees, land, and other
miscellaneous costs. The primary factor affecting capital cost is the amount of residuals produced,
which is dependent upon the design capacity of the water treatment plant and the treatment process
utilized (e.g., coagulation/filtration vs. lime softening).
The amount of waste generated plays a significant role in determining the handling and
disposal method to be utilized. Many handling methods which are suitable for smaller systems are
impractical for larger systems because of the significant land requirements. For larger systems that
process residuals on-site (as opposed to direct or indirect discharge), mechanical methods are
typically used because of the limited land requirements.
Operations and maintenance costs include labor, transportation, process materials and
chemicals, and maintenance. Many handling and disposal methods require extensive oversight
which can be a burden on small water systems. Generally, labor intensive technologies are more
suitable to large water systems. Transportation can also play a significant role in determining
appropriate handling and disposal options. If off-site disposal requires extensive transportation,
alternative disposal methods should be evaluated. Complex handling and disposal methods usually
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require more maintenance.
4.1.2 Methods for Estimating Residuals Handling and Disposal Costs
Residuals handling and disposal costs can be difficult to estimate. There are a number of
factors which affect capital and O&M costs, and disposal costs can be largely regional. EPA has
published two manuals for estimating residuals handling and disposal costs; Small Water System
Byproducts Treatment and Disposal Cost Document (DPRA, 1993a), and Water System Byproducts
Treatment and Disposal Cost Document (DPRA, 1993b). Both present a variety of handling and
disposal options, applications and limitations of those technologies;, and capital and O&M cost
equations.
4.2 RESIDUALS HANDLING OPTIONS
4.2.1 Gravity Thickening
Gravity thickening increases the solids content of filter backwash, sedimentation basins and
treatment process sludges. It is generally used as a pre-treatment for mechanical dewatering
processes, evaporation ponds, and storage lagoons.
Filter backwash streams are high volume, low solids slurries generated during the cleaning
of granular filter media. Backwash volume depends upon the number of filters and cleaning
frequency. Typical volumes range from 0.5 to 5 percent of the processed water flow with larger
plants creating less backwash per million gallons produced than small systems due to increased plant
efficiency (DPRA, 1993a). Backwash waters have an average solids concentration of 0.8 percent,
compared to coagulation sludges which are typically 0.5 to 2.0 percent (DPRA, 1993a).
When possible, backwash waters are recycled to the treatment process. In gravity thickening,
backwash waters are fed to a tank where settling occurs naturally. Sludges are discharged and
further treated for ultimate disposal, and the decant is either recycled or
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by other mechanical or non-mechanical dewatering processes. When backwash slurries cannot be
recycled or discharged to a surface water or POTW, they must be treated and disposed.
4.2.2 Mechanical Dewatering
Mechanical dewatering processes include centrifuges, vacuum-assisted dewatering beds, belt
filter presses, and plate and frame filter presses (DPRA, 1993a). Such processes generally have high
capital, as well as high O&M costs, compared to similar capacity non-mechanical dewatering
processes, e.g., storage lagoons. Due to the high costs such processes are generally not suitable for
application at very small water systems.
Filter presses have been used in industrial processes for years, and their use has been
increasing in the water treatment industry over the past several years. These devices have been
successfully applied to both lime and alum sludges. Prior to pressure filtration, alum sludges may
require the addition of lime to lower the resistance of the sludge to filtration. This is generally done
by adjusting the pH to approximately 11. Pre-conditioning also increases the sludge volume by as
much as 20 to 30 percent. Lime sludges can attain final solids concentrations of 40 to 70 percent,
while alum sludges may reach 35 to 50 percent total solids. Filter presses require little land, have
high capital costs, and are labor intensive (DPRA, 1993a). Capital and O&M costs are generally
higher than comparable non-mechanical dewatering alternatives. As a result, pressure filtration is
most applicable to larger water systems.
Centrifuges have also been used in the water industry for years. They are capable of
producing alum sludges with final solids concentrations of 15 to 30 percent and lime sludges with
65 to 70 percent total solids, based upon an influent solids concentration of 1 to 10 percent.
Centrifugation is a continuous process requiring minimal time (8 to 12 minutes) to achieve the
optimal sludge solids concentration. Centrifuges have low land requirements and high capital costs.
They are more labor intensive than non-mechanical alternatives, but less intensive than filter presses.
Again, due to the capital and O&M requirements centrifuges are more suitable for larger water
systems.
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4.2.3 Evaporation Ponds and Drying Beds
Evaporation ponds and drying beds are non-mechanical dewatering technologies wherein
favorable climatic conditions are used to dewater waste brines generated by treatment processes such
as reverse osmosis and ion exchange (DPRA, 1993a). Brine waste is discharged to a pond for
storage and evaporation. Ponds and drying beds are not generally suitable for alum and lime
sludges. Typically, such ponds are designed with large surface areas to allow the sun and wind to
effectively evaporate residual water. Size is determined by waste flow and storage capacity
requirements.
Evaporation ponds and drying beds are used primary for brine wastes generated by reverse
osmosis and ion exchange processes. Such processes produce large volumes of high TDS waste
streams and make mechanical dewatering processes, such as filter presses, impractical. Depending
upon the solids concentration of the brine waste stream, intermittent removal of solids may be
required. For brines with a total dissolved solid (TDS) content ranging from 15,000 to 35,000 mg/L,
solids will accumulate in the pond at a rate of 1A to 1V* inches per year (DPRA, 1993a). When the
depth of the solids reaches a predetermined level, flow to the pond is halted and evaporation
continues until the solids concentration is suitable for disposal.
Evaporation is an extremely land intensive handling option requiring shallow basins with
large surface areas. This can be an important consideration in densely populated regions. Reverse
osmosis produces a very large volume reject stream which increases the land requirement and
ultimately construction costs. As a result, evaporation ponds may not be suitable for large water
systems utilizing reverse osmosis. Evaporation ponds and drying beds have few operations and
maintenance requirements, but are only feasible in regions with favorable climatic conditions, i.e.,
high temperatures, low humidity, and low precipitation (DPRA, 1993a). Waste streams with low
TDS concentrations can allow a pond to operate for several years before solids accumulation
warrants removal.
4.2.4 Storage Lagoons
Lagoons are the most common, and often least expensive, method to thicken or dewater
treatment sludges; however, they are land intensive (DPRA, 1993a). Lagoons are lined ponds
4-4
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designed to collect and dewater sludge for a predetermined period of time. Dewatering occurs by
evaporation and decanting of the supernatant. Lagoon size is determined by the volume of sludge
produced and the storage time desired. As with evaporation ponds, when a lagoon reaches the
design capacity solids can be removed with heavy equipment and shipped for disposal.
Storage lagoons are best suited for dewatering lime softening process sludges, though they
have been applied with some success to coagulation/filtration process sludges. They can operate
under a variety of sludge flows and solids concentrations, and do not require chemical conditioning
of alum sludges (DPRA, 1993a). Typically, lime sludges enter the lagoon at three percent solids,
and can be dewatered to 50 to 60 percent solids, whereas alum sludges enter at one percent solids
and can be dewatered to 7 to 15 percent solids (DPRA, 1993a). Alum sludges do not typically
dewater well in storage lagoons. When the top layer of sludge is allowed to dry, it hardens, sealing
moisture in the layers below. Even after several years, alum sludges may require additional
dewatering to achieve the 20 percent solids content required at most landfills (DPRA, 1993a).
Further, thickened alum sludges can be difficult to remove from lagoons, and often require dredging
or vacuum pumping by knowledgeable operators.
As previously stated, lagooning is a land intensive process with limited applicability in
densely populated areas, or areas with limited land availability. Such areas need to compare the cost
of regular lagoon cleaning and disposal with land acquisition costs. Lagoons are best suited for areas
with favorable climatic conditions, i.e., high temperatures, low humidity, and low precipitation. In
northern climates, winter freezing can dehydrate alum sludges.
43 DISPOSAL ALTERNATIVES
4.3.1 Direct Discharge
Direct discharge to a surface water is a common method of disposal for water treatment
byproducts. No pretreatment or concentration of the byproduct stream is necessary prior to
discharge, and the receiving water dilutes the waste concentration and gradually incorporates the
sludge or brine (DPRA, 1993a).
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Discharge of liquid residuals containing arsenic to a surface water will be subject to
compliance with the National Pollution Discharge Elimination System (NPDES). NPDES
establishes limits based upon a variety of factors, including ambient contaminant levels, low flow
condition of the receiving water, and design flow of the proposed discharge. Most NPDES limits
for solids discharge are around 30 mg/L (AWWARF, 1998).
EPA has established water quality criteria under authority of the Clean Water Act. For
waters used for fish consumption the ambient water quality criterion for arsenic was set at 0.14 //g/L.
If the water source is used for drinking as well, that limit is reduced to 0.0175 /ug/L. These critera
will be used by state regulatory agencies to determine discharge limitations for arsenic depending
upon the classification of the receiving water. The allowable discharge is therefore affected by the
ability of the receiving water to assimilate the arsenic without exceeding the water quality criteria.
The primary cost associated with direct discharge is that of the piping. Accommodations
must be made for washout ports to prevent clogging because of sedimentation in pipelines. Valving
is necessary to control waste flow in the event of pipe bursts, and pipe must be laid at a sufficient
depth to prevent freezing in winter months. Direct discharge requires little oversight, and operator
experience and maintenance requirements are minimal. This method has been used to successfully
dispose of alum and lime sludges, as well as brine streams generated at reverse osmosis and ion
exchange water plants (DPRA, 1993a).
4.3.2 Indirect Discharge
In some cases, water treatment process sludges, slurries and brines may be discharged to a
POTW. This most often occurs when the treatment plant and POTW are under the same
management authority. This may require addition of a conveyance system to access the sanitary
sewer if an adequate system is not already in place (DPRA, 1993a).
Indirect discharge is a commonly used method of disposal for filter backwash and brine
waste streams. Coagulation/filtration and lime softening sludges have also been successfully
disposed of in this manner. However, the POTW must be able to handle the increased hydraulic and
solids loading. The capacity of the sewer system must also be considered when selecting indirect
discharge as a disposal option.
4-6
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The residuals generated from an arsenic treatment process will be classified as an industrial
waste since it contains contaminants, namely arsenic, which may impact the POTW. As a result,
discharge to a POTW is only acceptable when arsenic concentrations fall within the established
Technically Based Local Limits (TBLL) of the current Industrial Pretreatment Program (AWWARF,
1998). The Industrial Pretreatment Program serves to prevent NPDES violations, as well as
unacceptable accumulation of contaminants in POTW sludges and biosolids. TBLLs are
individually determined for each POTW, and take into account background levels of contamination
in the municipal wastewater. TBLLs for arsenic will typically be limited by the contamination of
biosolids rather than effluent limitations or process inhibition (AWWARF, 1998).
40 CFR 503 specifies the allowable limits for arsenic concentration in biosolids as a function
of disposal method. POTWs utilizing land application are subject to the Land Disposal Limit, Land
Application Ceiling Limit, and Land Application Clean Sludge Limit which are 73 mg/kg, 75 mg/kg
and 41 mg/kg, respectively. If the arsenic concentration exceeds the Clean Sludge criteria, land
application is limited to 41 kg per hectare (36.6 Ib/acre). As a result, most TBLLs are based upon
the Clean Sludge criterion. The typical POTW removal efficiency for arsenic is approximately 45
percent. Assuming biosolids production is around 1,200 pounds per million gallons of water treated,
the maximum allowable headworks loading will be around 0.109 pounds of arsenic per million
gallons of wastewater treated. This equates to a total (municipal and industrial) influent
concentration of around 13 /ug/L (AWWARF, 1998). As a result, if a water system has a background
arsenic concentration near 13 //g/L, it may not be possible to discharge to the sanitary sewer.
The primary cost associated with indirect discharge is that of the piping. Accommodations
must also be made for washout ports to prevent clogging because of sedimentation in pipelines.
Valving is necessary to control waste flow in the event of pipe bursts, and pipe must be laid at a
sufficient depth to prevent freezing in winter months. Additional costs associated with indirect
discharge may include lift stations, additional piping for access to the sewer system, or other
surcharges to accommodate the increased demands on the POTW.
4-7
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4.3.3 Dewatered Sludge Land Application
Dewatered sludge can be disposed by spreading the material over an approved land surface.
Application is dependent upon several variables, including soil and sludge chemistry or the crop
planted in the application field. Dewatered sludges are typically stored on site until they are
transported for application. Monitoring of soils, run off from land, application, and potentially
affected water sources is advisable to protect open land that may become cropland and to protect
local water quality (DPRA, 1993a).
As discussed in the previous section, land application of water treatment residuals containing
arsenic is limited to 41 mg/kg. If these concentrations cannot be achieved application of sludges is
limited to 41 kg arsenic per hectare. Due to the possibility of arsenic absorption by vegetation, non-
food chain fields are preferred for application. Land application is also limited by the availability
of land. In areas where grassland, farmland or forested land is unavailable, transportation can
significantly affect the cost effectiveness of this disposal option.
Land application can be a means of final disposal of lime softening, and to a lesser degree
coagulation/filtration, sludges. Lime sludges can be used in farmland to neutralize soil pH in place
of other commercial products. Alum sludges offer no benefit to soil chemistry and are generally
used as fill material.
4.3.4 Sanitary Landfill Disposal
Two forms of sanitary landfill are commonly used for disposal of water treatment
byproducts: monofills and commercial nonhazardous waste landfills (DPRA, 1993a). Monofills
only accept one type of waste, for example, fly ash or water treatment sludges. Commercial
nonhazardous waste landfills accept a variety of commercial and industrial wastes.
Sanitary landfills are regulated by both state and federal regulations. States have guidelines
on what types of waste can be landfilled, and determine construction and operation criteria. In many
cases, state requirements are more stringent than the federal regulations promulgated under the
Resource Conservation and Recovery Act (RCRA). The federal requirements include restrictions
on location, operation and design criteria, ground water monitoring requirements, corrective action
requirements, closure and post-closure requirements, and financial assurance.
4-8
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Landfill disposal requires that residuals be in a solid form and contain no free liquids.
Sanitary landfill disposal also requires that sludges meet specific criteria that determine if a waste
is hazardous. 40 CFR 261 establishes four characteristics of hazardous waste: flammability,
corrosivity, reactivity and toxicity. A waste must meet only one of the criteria to be considered
hazardous. With treatment residuals containing arsenic, toxicity is the primary characteristic of
concern.
EPA has established an analytical method, the Toxicity Characteristic Leaching Procedure
(TCLP), to measure the toxicity of a waste. The current TCLP limit for arsenic is 5 mg/L, which
is 100 times the current MCL of 50 /zg/L. If the MCL is lowered in the future, the TCLP value will
be lowered accordingly. For example, if the MCL were lowered to 20 /ug/L or 2 Mg/L, the TCLP
would be lowered to 2.0 mg/L or 200 ^g/L, respectively. As a result, water treatment residuals
containing arsenic may meet current sanitary landfill disposal criteria, but may not under a nature
regulatory framework.
Many water treatment facilities currently dispose of their waste in commercial or public-
owned landfills (DPRA, 1993a). In some parts of the country, decreasing landfill availability, rising
costs, and increasing regulations are making landfill disposal more expensive. As a result, the
benefits of monofills are being discussed within the industry. Costs associated with development
of monofills are generally less than those of a sanitary landfill (DPRA, 1993a). Monofills control
the type of waste disposed more strictly and limit the potential future liabilities, as well.
4.3.5 Hazardous Waste Landfill Disposal
Water treatment residuals containing arsenic which fail the TCLP test for toxicity must be
disposed in a designated and licensed hazardous waste landfill. Hazardous waste landfills are
regulated by the federal government under authority of RCRA or by individual states who have
received authorization under RCRA. Hazardous waste landfills are required to be permitted in
accordance with 40 CFR 270 which specifies landfill construction and operation criteria, and are
designed to isolate hazardous contaminants from the environment.
The primary limitation affecting disposal of arsenic containing residuals in a hazardous waste
landfill is the presence of free liquids. If any water treatment sludge contains free liquids, usually
4-9
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determined by the Paint Filter Liquids Test (SW-846, Method 9095), it is not suitable for landfilling.
Sludges containing free liquids must be stabilized or treated by another method to remove free
liquids prior to disposal in a hazardous waste landfill.
These facilities have extensive monitoring and operational requirements which cause the cost
of this method of disposal to be much greater than that of a typical sanitary landfill (AWWARF,
1998). If the residuals are determined to be hazardous, transportation to the landfill must be
manifested and the owner may never be free of responsibility for that waste. As a result, production
of a hazardous arsenic residual should be avoided if at all possible. Hazardous waste landfill
disposal is the most expensive disposal alternative discussed in this document, and should be used
only after all other disposal options have been exhausted.
4.4 RESIDUALS CHARACTERISTICS
4.4.1 Coagulation/Filtration
Coagulation/filtration (C/F) residual production is a function of coagulant type and
suspended solids content. For alum coagulation, approximately 0.26 pound of solids are produced
for every pound of alum added. For ferric coagulation, approximately 0.54 pound of solids are
produced for each pound of ferric chloride added (AWWARF, 1998). Sludge production is also
affected by the suspended solids content of the raw water.
Sludges removed from C/F sedimentation basins are high in water content and typically have
a solids content of less than 1.0 percent (AWWARF, 1998). As a result, such sludges are usually
discharged to a sanitary sewer or dewatered by one of the methods discussed earlier in this chapter.
Discharge to sanitary sewers is generally only an option for treatment plants with an average flow
of less than 10 million gallons per day.
Filter backwash is a high volume liquid waste stream with a solids content generally less than
1.0 percent. Typical volumes range from 1.0 to 2.0 percent of the treated flow. Backwash streams
are typically discharged to a sanitary sewer or processed using one of the mechanical methods
discussed in this chapter. As with sedimentation sludges, discharge of filter backwash streams to
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a sanitary sewer is generally only an option for treatment plants with an average flow of less than
10 million gallons per day.
The concentration of arsenic in the sedimentation sludges is a function of the amount of
coagulant used and the removal efficiency of the process. Using the design criteria specified in
Chapter 3, consider a treatment plant that utilizes ferric chloride at a dose of 25 mg/L. Neglecting
the impact of suspended solids, approximately 110 pounds of solids are produced for each million
gallons of water treated. The sedimentation sludges, or blowdown, will have a solids concentration
of 1.0 percent and be produced at a rate of 1,400 gallons per million gallons of water treated.
Assuming 40 yUg/L of arsenic is removed from the influent during the treatment proces, the result
is 0.33 pounds of arsenic per million gallons of water treated. The resulting residuals arsenic
concentration is 28.0 mg/L, or 3,000 mg/kg on a dry weight basis. Filter backwash will be
produced at a rate of 10,000 gallons per million gallons of water produced, and will have an
approximate solids concentration of 1 percent (10,000 mg/L).
C/F blowdown and filter backwash are high volume, low solids content waste streams.
Gravity thickening may be used as a pretreatment for C/F sludges and backwash prior to handling
by other mechanical or non-mechanical dewatering processes. Filter presses are capable of attaining
final solids contents in the range of 35 to 50 percent, while scroll centrifuges may achieve final
solids contents of 15 to 30 percent. Evaporation ponds and storage lagoons may be suitable for
smaller treatment plants, but because they are land intensive may not be applicable for large water
systems.
Disposal of C/F arsenic residuals is largely dependent upon influent arsenic concentration,
coagulant dose and suspended solids content. Disposal by direct discharge to a surface water is not
likely. In the example presented earlier in this section, the typical solids concentration of 1 .0
percent, or 10,000 mg/L, far exceeds the usual NPDES limit of around 30 mg/L.
The blowdown arsenic concentration in the above example is 28,000 ,ug/L. Most Industrial
Protection Programs have TBLLs in the range of 50 //g/L to 1,000 yUg/L (AWWARF, 1998). As a
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result, it is unlikely that indirect discharge to a POTW is an acceptable disposal method for C/F sludges.
Depending upon the arsenic concentration of C/F sludges, land application may be a suitable
disposal method. Total arsenic should not exceed 41 mg/kg if sludges are to be applied with no
restrictions. Sludges with higher arsenic concentration may be land applied providing the total
loading does not exceed 41 kg per hectare. In the previous example, the arsenic concentration is
2,930 mg/kg on a dry weight basis. As a result, this sludge may be hind applied, but loading must
not exceed 41 kg per hectare.
All C/F sludges must be dewatered prior to landfill disposal. If the residuals pass the TCLP
test they may be disposed in a sanitary landfill. Otherwise, residuals must be disposed in a
hazardous waste landfill. Tests conducted by the University of Colorado indicate that most C/F
sludges will pass the TCLP test (AWWARF, 1998). Hazardous waste landfill disposal should only
be used as a last resort if waste fails the TCLP test.
4.4.2 Enhanced Coagulation
Enhanced coagulation is a modified C/F process that includes increased coagulant dosage,
reduction in process pH, or both. As a result, enhanced coagulation process residuals are nearly
identical to typical C/F residuals. The exception is increased solids production as a result of the
increased coagulant dosage.
Sludges removed from enhanced coagulation sedimentation basins are high in water content
and typically have a solids content of approximately 1.0 percent (AWWARF, 1998). As a result,
such sludges are usually discharged to a sanitary sewer or dewatered by one of the methods
discussed earlier in this chapter. Discharge to sanitary sewers is generally only an option for
treatment plants with an average flow of less than 10 million gallons per day.
Filter backwash is a high volume liquid waste stream with a solids content generally less than
1.0 percent Typical volumes range from 1.0 to 2.0 percent of the treated flow. Backwash streams
are typically discharged to a sanitary sewer or processed using one of the mechanical methods
discussed in this chapter. As with sedimentation sludges, discharge of filter backwash streams to
a sanitary sewer is generally only an option for treatment plants with an average flow of less than
10 million gallons per day.
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The concentration of arsenic in the treatment process residuals is a function of the amount
of coagulant used and the removal efficiency of the process. Using the design criteria specified in
Chapter 3, consider a treatment plant that utilizes ferric chloride at a dose of 35 mg/L. Neglecting
the impact of suspended solids, approximately 160 pounds of solids are produced for each million
gallons of water treated. The sedimentation sludges, or blowdown, will have a solids concentration
of 1.0 percent and be produced at a rate of 1,900 gallons per million gallons of water treated.
Assuming 40 Mg/L of arsenic is removed from the influent during the treatment process, the result
is 0.33 pounds of arsenic per million gallons of water treated. The resulting residuals arsenic
concentration is 21 mg/L, or 2,060 mg/kg on a dry weight basis. Filter backwash will be
produced at a rate of 10,000 gallons per million gallons of water produced, and will have an
approximate solids concentration of 1 percent (10,000 mg/L).
Enhanced coagulation blowdown is a high volume, low solids content waste stream. Typical
solids contents range from 0.5 to 2.0 percent, depending upon the coagulant type. Gravity
thickening may be used as a pretreatment for C/F sludges prior to handling by other mechanical or
non-mechanical dewatering processes. Filter presses are capable of attaining final sludge solids
contents in the range of 35 to 50 percent, while scroll centrifuges may achieve final solids contents
of 15 to 30 percent. Evaporation ponds and storage lagoons may be suitable for smaller treatment
plants, but because they are land intensive may not be applicable for large water systems.
Disposal of enhanced coagulation arsenic residuals is largely dependent upon influent arsenic
concentration, coagulant dose, and suspended solids content. Disposal by direct discharge to a
surface water is not likely. In the example presented earlier in this section, the solids concentration
is 1.0 percent, or 10,000 mg/L. This is a typical solids content for C/F sludges, and far exceeds the
usual NPDES limit of around 30 mg/L.
The blowdown arsenic concentration in the above example is 21,000 ,ug/L. Most Industrial
Protection Programs have TBLLs in the range of 50 /ug/L to 1,000 //g/L (AWWARF, 1998). As a
result, it is unlikely that indirect discharge to a POTW is an acceptable disposal method for C/F
sludges.
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Depending upon the arsenic concentration of C/F sludges, land application may be a suitable
disposal method. Total arsenic should not exceed 41 mg/kg if sludges are to be applied with no
restrictions. Sludges with higher arsenic concentration may be land applied providing the total
loading does not exceed 41 kg per hectare. In the previous example, the arsenic concentration is
2,060 mg/kg on a dry weight basis. As a result, this sludge may be Land-applied, but loading must
not exceed 41 kg per hectare.
All enhanced coagulation sludges must be dewatered prior to landfill disposal. If the
residuals pass the TCLP test they may be disposed in a sanitary landfill. Otherwise, residuals must
be disposed in a hazardous waste landfill. Tests conducted by the University of Colorado indicate
that enhanced coagulation sludges will pass the TCLP test (AWWARP, 1998). Hazardous waste
landfill disposal should only be used as a last resort if waste fails the TCLP test.
4.4.3 Direct Filtration
Direct filtration is a modified C/F process that lacks the sedimentation unit process.
Accordingly, direct filtration residuals are the result of filter backwash, and typically have lower
TDS concentrations than a typical C/F process. This is due to the reduced coagulant dose. Sludge
production is also affected by the suspended solids content of the raw water.
Backwash from direct filtration plants is high in water content and typically has a solids
content of less than 1.0 percent. As a result, such sludges are usually discharged to a sanitary sewer
or dewatered by one of the methods discussed earlier in this chapter. Discharge to sanitary sewers
is generally only an option for treatment plants with an average daily flow of less than 10 million
gallons per day.
The concentration of arsenic in the treatment process residuals is a function of the amount
of coagulant used and the removal efficiency of the process. Using the design criteria specified in
Chapter 3, consider a treatment plant that utilizes ferric chloride at a dose of 10 mg/L. Neglecting
the impact of suspended solids, approximately 45 pounds of solids are produced for each million
gallons of water treated. The filter backwash will have a solids concentration of 0.0005 percent and
be produced at a rate of 20,000 gallons per million gallons of water treated. Assuming 40 /^g/L of
arsenic is removed from the influent during the treatment process, the result is 0.33 pounds of
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arsenic per million gallons of water treated. The backwash residuals arsenic concentration is
approximately 2.0 mg/L, or 7,333 mg/kg on a dry weight basis.
Direct filtration backwash is a high volume, low solids content waste stream. Gravity
thickening may be used as a pretreatment for sludges prior to handling by other mechanical or non-
mechanical dewatering processes. Filter presses are capable of attaining final sludge solids contents
in the range of 35 to 50 percent, while scroll centrifuges may achieve final solids contents of 15 to
30 percent. Evaporation ponds and storage lagoons may be suitable for smaller treatment plants, but
because they are land intensive may not be applicable for large water systems.
Disposal of direct filtration arsenic residuals is largely dependent upon influent arsenic
concentration, coagulant dose and suspended solids content. Disposal by direct discharge to a
surface water may be possible. In the example presented earlier in this section, the typical solids
concentration is 0.0005 percent, or 5 mg/L, which meets the usual NPDES limit of around 30 mg/L.
However, the above example neglects suspended solids and is based upon a backwash volume of
20,000 gallons. Higher influent suspended solids and/or smaller backwash volumes will impact the
ability of a treatment facility to dispose of direct filtration residuals by direct discharge.
The backwash arsenic concentration in the above example is 2,000 /ug/L. Most Industrial
Protection Programs have TBLLs in the range of 50 jzg/L to 1,000 jtg/L (AWWARF, 1998). As a
result, it is unlikely that indirect discharge to a POTW is an acceptable disposal method for C/F
sludges.
Depending upon the arsenic concentration of direct filtration sludges, land application may
be a suitable disposal method. Total arsenic should not exceed 41 mg/kg if sludges are to be applied
with no restrictions. Sludges with higher arsenic concentration may be land applied providing the
total loading does not exceed 41 kg per hectare. In the previous example, the arsenic concentration
is 7,333 mg/kg on a dry weight basis. As a result, this sludge may be land applied, but loading must
not exceed 41 kg per hectare.
All sludges must be dewatered prior to landfill disposal. If the residuals pass the TCLP test
they may be disposed in a sanitary landfill. Otherwise, residuals must be disposed in a hazardous
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waste landfill. Tests conducted by the University of Colorado indicate that direct filtration sludges
will pass the TCLP test (AWWARF, 1998). Hazardous waste landfill disposal should only be used
as a last resort if waste fails the TCLP test.
4.4.4 Coagulation Assisted Microfiltration
Coagulation assisted microfiltration is a modified C/F process wherein the
flocculation/sedimentation and filtration unit processes are replaced by microfiltration. Residuals
generated by this process consist of a filter backwash stream containing a dilute Fe(OH)3 precipitate
concentration. Based upon the design criteria in Chapter 3, assume: a ferric chloride dose of 10
mg/L, a recovery rate of 95 percent, and that each pound of ferric chloride added produces 0.54
pounds of precipitate. Under these conditions, 52,600 gallons of backwash containing 45 pounds
of precipitate at a concentration of 103 mg/L will be produced for every million gallons of water
produced. Assuming an arsenic removal of 40 ,ug/L, the backwash will have an approximate arsenic
concentration of 744 ,ug/L, or 7,333 mg/kg on a dry weight basis.
Residuals from coagulation assisted microfiltration processes will be a very dilute slurry with
a solids concentration of approximately 0.01% (103 mg/L). Gravity thickening may be used as a
pre-treatment for other mechanical or non-mechanical dewatering options. Filter presses and
centrifuges are appropriate methods of residuals handling. However, these methods are capital
intensive and may not be appropriate for extremely large systems. Evaporation ponds and storage
lagoons are also appropriate for coagulation assisted microfiltration residuals handling. Both require
little oversight and maintenance, but are land intensive. As such, these may not be appropriate for
large systems. A thorough comparison of handling options should be conducted to select the most
cost effective method.
Direct discharge of coagulation assisted microfiltration residuals is not a likely disposal
option. In the example above, the solids content of the residuals stream is 103 mg/L. This is greater
than the typical NPDES limit of 30 mg/L. As a result, direct discharge is an unlikely disposal option
for coagulation assisted microfiltration residuals.
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The effluent arsenic concentration of 744 /zg/L estimated for coagulation assisted
microfiltration facilities is within the range of TBLLs (50 //g/L to 1,000 //g/L). Therefore, discharge
to a sanitary sewer will be an acceptable method of disposal for coagulation assisted microfiltration
residuals containing arsenic. The Industrial Protection Program should be checked prior to disposal
to verify that the arsenic concentration of the residuals does not exceed the TBLL.
Land application may be a suitable disposal method for coagulation assisted microfiltration
sludges. The predicted arsenic concentration of 7,333 mg/kg is much higher than the Clean Sludge
criteria of 41 mg/kg, as well as the Land Application Ceiling limit of 75 mg/kg. Coagulation
assisted microfiltration sludges may be land applied provided the total loading does not exceed 41
kg of arsenic per hectare.
All coagulation assisted microfiltration sludges must be dewatered prior to landfill disposal.
If the residuals pass the TCLP test they may be disposed in a sanitary landfill. Otherwise, residuals
must be disposed in a hazardous waste landfill. Tests conducted by the University of Colorado
indicate that sludges should pass the TCLP test (AWWARF, 1998). Hazardous waste landfill
disposal should only be used as a last resort if waste fails the TCLP test.
4.4.5 Lime Softening
The quantity of residuals produced at lime softening (LS) facilities is typically much greater
than the quantity produced by C/F plants (AWWARF, 1998). The quantity of sludges produced is
a function of water hardness. LS for carbonate hardness removal produces approximately twice the
amount of solids per pound of hardness removed than non-carbonate hardness removal.
LS plants typically produce 1,000 to 8,000 pounds of solid per million gallons of water
treated depending upon the hardness of the water (AWWARF, 1998). Arsenic concentrations,
however, are generally lower than C/F sludges due to the increased volume of solids produced.
Using the design criteria hi Chapter 3, assume a treatment plant generates 20,000 gallons of
blowdown per million gallons of treated water. A solids concentration of 1.0 percent will result in
1,665 pounds of solids per million gallons of treated water. If 40 //g/L of arsenic are removed in the
process, 0.33 pounds of arsenic are produced per million gallons treated. This equates to 2.0 mg/L
in the blowdown, or 200 mg/kg on a dry weight basis.
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Selection of Handling ami-Disposal Options
LS blowdown is slightly denser than C/F blowdown. Typical solids contents range from 1.0
to 4.0 percent, depending upon the raw water hardness. Gravity thickening may be used as a
pretreatment for LS sludges prior to handling by other mechanical or non-mechanical dewatering
processes. Filter presses are capable of attaining final LS sludge solids contents in the range of 40
to 70 percent, while scroll centrifuges may achieve final solids contents of 65 to 70 percent.
Evaporation ponds and storage lagoons may be suitable for smaller treatment plants, but because
they are land intensive may not be applicable for large water systems.
Direct discharge of LS sludges to a surface water is not a likely disposal alternative. LS
sludges are typically 1.0 to 4.0 percent solids (10,000 to 40,000 mg/L) and will exceed the 30 mg/L
limit in most NPDES permits. Further, discharge to a sanitary sewer is not appropriate. In the
example presented earlier in this section, the arsenic concentration of the blowdown is approximately
2.0 mg/L which exceeds the typical TBLL (50 to 1,000 Mg/L) of an Industrial Protection Program.
Land application of LS treatment sludges is one possible disposal alternative. Based upon
the above example, it appears that LS sludges will exceed the Clean Sludge criteria for arsenic of
41 mg/kg. Application would therefore be limited to 41 kg of arsenic per hectare.
LS sludges will require dewatering prior to landfill disposal. If the residuals pass the TCLP
test they may be disposed in a sanitary landfill. Otherwise, residuals must be disposed in a
hazardous waste landfill. Tests conducted by the University of Colorado indicate that LS sludges
will pass the TCLP test (AWWARF, 1998). Hazardous waste landfill (disposal should only be used
as a last resort if waste fails the TCLP test.
4.4.6 Enhanced Lime Softening
Enhanced LS is a modified LS process wherein lime dosage is increased. Residuals produced
are similar to those of a typical LS treatment process. The quantity of sludges produced is a function
of water hardness. Enhanced LS for carbonate hardness removal produces approximately twice the
amount of solids per pound of hardness removed than non-carbonate hardness removal.
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Enhanced LS plants typically produce 1,000 to 8,000 pounds of solid per million gallons of
water treated depending upon the hardness of the water (AWWARF, 1998). Arsenic concentrations,
however, are generally lower than C/F sludges due to the increased volume of solids produced.
Using the design criteria in Chapter 3, assume a treatment plant generates 20,000 gallons of
blowdown per million gallons of treated water. A solids concentration of 1 .0 percent will result in
1,665 pounds of solids per million gallons of treated water. If 40 /ug/L of arsenic are removed in the
process, 0.33 pounds of arsenic are produced per million gallons treated. This equates to 2.0 mg/L
in the blowdown, or 200 mg/kg on a dry weight basis.
Typical enhanced LS blowdown solids contents range from 1.0 to 4.0 percent, depending
upon the raw water hardness. Gravity thickening may be used as a pretreatment for sludges prior
to handling by other mechanical or non-mechanical dewatering processes. Filter presses are capable
of attaining final sludge solids contents in the range of 40 to 70 percent, while scroll centrifuges may
achieve final solids contents of 65 to 70 percent. Evaporation ponds and storage lagoons may be
suitable for smaller treatment plants, but because they are land intensive may not be applicable for
large water systems.
Direct discharge of enhanced LS sludges to a surface water is not a likely disposal
alternative. LS sludges are typically 1.0 to 4.0 percent solids (10,000 to 40,000 mg/L) and will
exceed the 30 mg/L limit in most NPDES permits. Further, discharge to a sanitary sewer is not
appropriate. In the example presented earlier in this section, the arsenic concentration of the
blowdown is approximately 2.0 mg/L which exceeds the typical TBLL (50 to 1 ,000 Mg/L) of an
Industrial Protection Program.
Land application of treatment sludges is one possible disposal alternative. Based upon the
above example, it appears that enhanced LS sludges will exceed the Clean Sludge criteria for arsenic
of 41 mg/kg. Application would therefore be limited to 41 kg of arsenic per hectare. As with C/F
sludges, once this loading is achieved it may never be used again for disposal purposes.
Sludges will require dewatering prior to landfill disposal. If the residuals pass the TCLP test
they may be disposed in a sanitary landfill. Otherwise, residuals must be disposed in a hazardous
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waste landfill. Tests conducted by the University of Colorado indicate that enhanced LS sludges will
pass the TCLP test (AWWARF, 1998). Hazardous waste landfill disposal should only be used as
a last resort if waste fails the TCLP test.
4.4.7 Ion Exchange
Ion exchange (DC) uses a synthetic resin in a chloride form for arsenic removal. With time,
the efficiency of the resin is reduced as exchange sites are depleted. The IX resin can be regenerated
using a NaCl solution. The regenerant is added at a rate of approximately 2 equivalents per
equivalent of resin, i.e., 15 pounds of salt per cubic foot of resin. Regeneration requires
approximately 2.8 BV of brine and 1.2 BV displacement rinse. Therefore, 4 to 5 BV of waste are
produced per regeneration cycle. (AWWARF, 1998).
Arsenic removal is severely affected by sulfate concentration of the source water. Assuming
the sulfate concentration limits the run length to 1,000 BV before regeneration is required, at an
influent arsenic concentration of 40 //g/L, approximately 1,132 mg of arsenic can be removed per
cubic foot of resin. Regeneration will produce a brine waste solution with an arsenic concentration
of around 10 mg/L and a brine concentration of about 20,000 mg/L (AWWARF, 1998). Arsenic can
be precipitated from the brine stream using ferric chloride. The resulting precipitate has an arsenic
concentration of approximately 14,250 mg/kg on a dry weight basis.
Selection nf Handling and Disposal Options
Evaporation ponds, drying beds, and storage lagoons are often used for brine waste stream
handling. In regions with favorable climatic conditions, evaporation ponds may be the preferred
handling option. Evaporation ponds are land intensive and require shallow basins with large surface
areas. As such, ponds may not be suitable to areas where available land is scarce or acquisition costs
cannot be justified for construction of a pond. If construction is feasible;, operations and maintenance
are minimal and make this a primary candidate for IX residuals handling.
Direct discharge of DC residuals to a receiving surface water is an unlikely disposal
alternative. Assume the brine waste stream generated has an arsenic concentration of 10 mg/L.
Now, assume the state regulatory agency has set an ambient arsenic water quality standard at 20.5
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,ug/L for fish consumption. IX will produce approximately 4,000 gallons of brine waste for each
million gallons of treated water. Based upon this scenario, a dilution factor of around 487 is needed
for the receiving water to assimilate the arsenic from the brine stream. In other words, a one million
gallon per day treatment plant would require a dilution flow of 3.1 cfs to assimilate the 4,000 gallons
of waste brine generated (AWWARF, 1998). Furthermore, to be discharged directly to a receiving
water the brine stream must pass the whole effluent toxicity (WET) test. It is unlikely that the waste
brine would pass the WET test due to the high arsenic concentration and salt content (AWWARF,
1998). Therefore, even if a stream had sufficient flow for dilution of the IX residuals, it is unlikely
discharge would be allowed.
The 10 mg/L concentration of the brine waste stream far exceeds the typical TBLL of 50 to
1,000 Mg/L- Therefore, it is unlikely that IX residuals could be discharged to a sanitary sewer for
treatment at a POTW. If the solids content of the waste brine were increased to 5 percent by ferric
chloride precipitation, the arsenic concentration is increased to 712 mg/L (AWWARF, 1998). Again,
this exceeds typical TBLL values and would not likely be a candidate for discharge to a sanitary
sewer.
Land application of IX brine streams is an unlikely disposal alternative. The high salinity
of the residual stream would result in a significant increase in the salt content of the receiving soil.
This salinity build-up would make plant growth virtually impossible. Therefore, an alternative
technology should be selected for disposal or IX residuals. If land application is considered
appropriate, arsenic loading may not exceed 41 kg per hectare. Once land has been used for arsenic
residuals disposal it may never be used for that purpose again.
IX residuals may be disposed at a sanitary or hazardous waste landfill, but will require
extensive dewatering. The solids content can be increased to approximately 5 percent by ferric
chloride precipitation. The precipitate will require dewatering by one of the methods presented in
this chapter prior to disposal. The precipitate will likely pass the TCLP test and be a candidate for
disposal at a sanitary landfill (AWWARF, 1998). However, should the residuals fail the TCLP test,
disposal at a hazardous waste landfill will be required.
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4.4.8 Activated Alumina
Activated alumina (AA) can be operated with or without regeneration. Small water systems
may also choose to operate without optimizing process pH. Systems optimizing pH and regenerating
will produce a regenerant waste solution with a pH of approximately 12 and high in dissolved solids,
aluminum, and arsenic (AWWARF, 1 998). Arsenic can be removed from the solution by aluminum
hydroxide precipitation and reducing the regenerant pH. Regeneration of AA is accomplished using
15 to 25 bed volumes (BV) of 2N NaOH, 7 BV of rinse, and 15 BV of 2N H2SO4 for neutralization.
Thus, the total volume of waste produced in approximately 42 BV pa: regeneration cycle. Systems
who choose not to regenerate will be required to dispose of the spent resin.
Assume that an AA system operates at optimal pH, regenerates and can treat 1 0,000 B V of
water before it reaches exhaustion, and that approximately 40 /ug/L of arsenic are being removed
from the source water. The resulting waste solution will have an approximate arsenic concentration
of 9.52 mg/L. Ferric chloride precipitation of the residual arsenic would produce a residual with an
arsenic concentration of around 14,250 mg/kg on a dry weight basis (AWWARF, 1998).
Assume a sytem with a raw water pH of 7 chooses not to optimize pH, operates on a "throw
away" basis, i.e., no regeneration, and removes approximately 40 [
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Direct discharge of AA residuals to a receiving surface water is an unlikely disposal
alternative. Assume the brine waste stream generated has an arsenic concentration of 9.52 mg/L.
Now, assume the state regulatory agency has set an ambient arsenic water quality standard at 20.5
^ig/L for fish consumption. AA will produce approximately 4,200 gallons of brine waste for each
million gallons of treated water. Based upon this scenario, a dilution factor of around 463 is needed
for the receiving water to assimilate the arsenic from the brine stream. In other words, a one million
gallon per day treatment plant would require a dilution flow of 3.0 cfs to assimilate the 4,200 gallons
of waste brine generated (AWWARF, 1998). Furthermore, to be discharged directly to a receiving
water the brine stream must pass the WET test. It is unlikely that the waste brine would pass the
WET test due to the high arsenic concentration (AWWARF, 1998). Therefore, even if a stream had
sufficient flow for dilution of the AA residuals, it is unlikely discharge would be allowed.
The 9.52 mg/L concentration of the brine waste stream far exceeds the typical TBLL of 50
to 1,000 /ug/L. Therefore, it is unlikely that AA residuals could be discharged to a sanitary sewer
for treatment at a POTW. If the solids content of the waste brine were increased to 5 percent by
ferric chloride precipitation, the arsenic concentration is increased to 712 mg/L (AWWARF, 1998).
Again, this exceeds typical TBLL values and would not likely be a candidate for discharge to a
sanitary sewer.
Land application of AA residuals may be inappropriate. The aluminum content of the waste
solution may combine with phosphorus in the soil matrix and prevent uptake of the phosphorus by
vegetation. If land application is considered appropriate, arsenic loading may not exceed 41 kg per
hectare. Once land has been used for arsenic residuals disposal it may never be used for that purpose
again.
AA residuals may be disposed at a sanitary or hazardous waste landfill, but will require
extensive dewatering. The solids content can be increased to approximately 5 percent by aluminum
hydroxide precipitation. The precipitate will require dewatering by one of the methods presented
in this chapter prior to disposal. Tests conducted at the University of Colorado indicate the
precipitate will likely pass the TCLP test and be a candidate for disposal at a sanitary landfill
(AWWARF, 1998). However, should the residuals fail the TCLP test, disposal at a hazardous waste
landfill will be required.
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Residuals from systems operating on a "throw away" basis, i.e., not regenerating, can
dispose of process residuals in a sanitary or hazardous waste landfill. Spent resin will have to pass
the TCLP test before disposal at a sanitary landfill is possible.
4.4.9 Microfiltration
Microfiltration (MF) membrane pore size is too large to nanove dissolved or colloidal
arsenic, but is capable of removing particulate arsenic. Considering particulate arsenic
concentrations in groundwater are generally less than 10 percent, ;and surface water particulate
arsenic concentrations vary from 0 to 70 percent, MF alone may not be a viable removal technology.
Recovery rates are much higher than RO and NF processes, with typical recovery approaching 99
percent. The reject stream contains elevated levels of arsenic and other contaminants that are
removed from the source water by the MF membranes.
As an example, assume that a MF system has a recovery rate of 95 percent, arsenic rejection
by the membrane is 20 percent, and the feed arsenic concentration is 50 /ug/L. The reject will have
an arsenic concentration of approximately 190 //g/L. This system will also produce 52,600 gallons
of reject for every million gallons of treated water. Due to the large volume of reject water, it may
not be feasible to implement ferric chloride precipitation to remove the arsenic from the reject
stream.
Selection of Handling and Disposal Options
MF generates waste streams with high suspended solids content. Evaporation ponds and
drying beds have been used successfully for similar waste streams in the past. Both are land
intensive handling options and may not be suitable for large MF facilities, or for water systems
where land acquisition costs make construction infeasible. If determined to be appropriate,
evaporation ponds and drying beds require little oversight and maintenance is minimal. Depending
upon solids concentrations in the reject stream, other mechanical and non-mechanical dewatering
devices may also be applicable.
Direct discharge of MF reject streams to a surface water may not be a feasible disposal
alternative. MF reject streams typically contain greater than 100 mg/L total solids, and will exceed
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the 30 mg/L limit in most NPDES permits. Discharge to a sanitary sewer may be appropriate. In
the example presented earlier in this section, the arsenic concentration of the reject is approximately
190 Mg/L which is within the lower bounds of the typical TBLL (50 to 1,000 /ug/L) of an Industrial
Protection Program.
Land application may be a suitable disposal method for MF reject streams. MF reject
streams may be land applied providing the total loading does not exceed 41 kg of arsenic per
hectare.
All MF sludges must be dewatered prior to landfill disposal. If the residuals pass the TCLP
test they may be disposed in a sanitary landfill. Otherwise, residuals must be disposed in a
hazardous waste landfill. Hazardous waste landfill disposal should only be used as a last resort if
waste fails the TCLP test.
4.4.10 Ultrafiltration
Ultrafiltration (UF) membranes are primarily used for removal of colloidal and particulate
contaminants. Considering arsenic found in groundwater is typically less than 10 percent particulate,
and surface waters contain 0 to 70 percent particulate arsenic, UF may not be a viable arsenic
removal technology. The recovery rate for UF is generally higher than that of RO or NF, with
typical recovery approaching 95%. The reject stream contains elevated levels of arsenic and other
contaminants that are removed from the source water by the UF membranes.
As an example, assume that a membrane system has a recovery rate of 90 percent, arsenic
rejection by the membrane is 70 percent, and the feed arsenic concentration is 50 yUg/L. The reject
will have an arsenic concentration of approximately 315 //g/L. This system will also produce
111,000 gallons of reject for every million gallons of treated water. Due to the large volume of
reject water, it may not be feasible to implement ferric chloride precipitation to remove the arsenic
from the reject stream.
Selection of Handling and Disposal Options
UF generates high volume, high suspended solids waste streams. Evaporation ponds and
drying beds have been used successfully for similar waste streams in the past. Both are land
4-25
-------
intensive handling options and may not be suitable for large UF facilities, or for water systems where
land acquisition costs make construction infeasible. If determined to be appropriate, evaporation
ponds and drying beds require little oversight and maintenance is minimal. Depending upon solids
concentrations in the reject stream, other mechanical and non-mechanical dewatering devices may
also be applicable.
Direct discharge of UF reject streams to a surface water may be likely, providing solids
concentrations do not exceed the 30 mg/L limit in most NPDES permits. Discharge to a sanitary
sewer may also be an appropriate method of disposal. In the example presented earlier in this
section, the arsenic concentration of the reject stream is approximately 660 /ug/L which is in the
upper bounds of the typical TBLL (50 to 1,000 /^g/L) of an Industrial Protection Program.
Land application is a possible candidate for disposal of UF arsenic containing residuals. The
reject water would have to be applied at a rate matching the evapotransportation requirements of the
cover crop grown and is limited to 41 kg of arsenic per hectare.
All UF sludges must be dewatered prior to landfill disposal. If the residuals pass the TCLP
test they may be disposed in a sanitary landfill. Otherwise, residuals must be disposed in a
hazardous waste landfill. Hazardous waste landfill disposal should only be used as a last resort if
waste fails the TCLP test.
4.4.11 Nanofiltration
NF membranes are primarily used for softening, removing the larger, divalent ions
associated with hardness (AWWARF, 1998). The recovery rate, or percent of water passing through
the membrane rather than being rejected by the membrane, is dependent upon the source water
quality. Typical recovery rates approach 85% percent. The reject stream contains elevated levels
of arsenic and other contaminants that are removed from the source water by the NF membranes.
Based upon the design criteria in Chapter 3, assume that a NF system has a recovery rate of
85 percent. Arsenic rejection by the membrane is 85 percent and the feed arsenic concentration is
50/zg/L. The reject will have an arsenic concentration of approximately 225//g/L. This system will
also produce 176,000 gallons of reject for every million gallons of treated water (AWWARF, 1998).
4-26
-------
Due to the large volume of reject water, it may not be feasible to implement ferric chloride
precipitation to remove the arsenic from the reject stream.
Selection of Handllng_andJ)ispojal_Qp.tiQjis
NF generates high volume, high TDS waste streams. Evaporation ponds and drying beds
have been used successfully for similar waste streams in the past. Both are land intensive handling
options and may not be suitable for large facilities, or for water systems where land acquisition costs
make construction infeasible. If determined to be appropriate, evaporation ponds and drying beds
require little oversight and maintenance is minimal. Because of the high volume of NF waste
streams, other mechanical and non-mechanical dewatering devices are generally not applicable.
Direct discharge of NF residuals to a receiving surface water is one possible disposal
alternative. For example, assume the brine waste stream generated has an arsenic concentration of
225 //g/L, and the state regulatory agency has set an ambient arsenic water quality standard at 20.5
Mg/L for fish consumption. NF will produce approximately 176,000 gallons of brine waste for each
million gallons of treated water. Based upon this scenario, a dilution factor of around 12 is needed
for the receiving water to assimilate the arsenic from the brine stream. In other words, a one million
gallon per day treatment plant would require a dilution flow of 3.1 cfs to assimilate the 176,000
gallons of waste brine generated. Furthermore, to be discharged directly to a receiving water the
brine stream must pass the WET test. Due to the low arsenic concentration in the NF process
effluent, it is unlikely that the waste brine would fail the WET test (AWWARF, 1998). Therefore,
direct discharge to a receiving water may be a possible disposal alternative for NF treatment
facilities.
The effluent arsenic concentration of 225 //g/L estimated for the NF process is well within
the bounds of the range of TBLLs (50 yug/L to 1,000 Mg/L)- Therefore, in most cases discharge to
a sanitary sewer will be an acceptable method of disposal for NF residuals containing arsenic. The
appropriate Industrial Protection Program should be checked prior to disposal to verify that the
arsenic concentration of the residuals does not exceed theTBLL.
NF treatment produces high volume liquid waste streams. As a result, land application is an
unlikely candidate for disposal of arsenic containing residuals. The reject water would have to be
4-27
-------
applied at a rate matching the evapotransportation requirements of the cover crop grown and is
limited to 41 kg of arsenic per hectare. Therefore, land application is deemed inappropriate for NF
residuals disposal (AWWARF, 1998).
The liquid waste produced by NF treatment can not be sent to a sanitary or hazardous waste
landfill. The high free liquids content makes dewatering uneconomical. -The arsenic concentration
is well below TCLP limits and makes hazardous waste landfill disposal unnecessary. Landfill
disposal of NF residuals is inappropriate (AWWARF, 1998).
4.4.12 Reverse Osmosis
RO membranes will remove much smaller ions typically associated with TDS (AWWARF,
1998). The recovery rate, or percent of water passing through the membrane rather than being
rejected by the membrane, is dependent upon the source water quality. Typical recovery rates vary
from 30 to 85 percent. The reject stream contains elevated levels of arsenic and other contaminants
that are removed from the source water by the RO membranes.
Based upon the design criteria in Chapter 3, assume that a RO membrane system has a
recovery rate of 85 percent. Arsenic rejection by the membrane is 95 percent and the feed arsenic
concentration is 50 /ag/L. The reject will have an arsenic concentration of approximately 265 //g/L.
This system will also produce 176,000 gallons of reject for every million gallons of treated water
(AWWARF, 1998). Due to the large volume of reject water, it may not be feasible to implement
ferric chloride precipitation to remove the arsenic from the reject stream.
Selection of Handling and Disposal Options
RO generates high volume, high TDS waste streams. Evaporation ponds and drying beds
have been used successfully for similar waste streams in the past. Both are land intensive handling
options and may not be suitable for large RO facilities, or for water systems where land acquisition
costs make construction infeasible. If determined to be appropriate, evaporation ponds and drying
beds require little oversight and maintenance is minimal. Because of the high volume of RO waste
streams, other mechanical and non-mechanical dewatering devices are generally not applicable.
4-28
-------
Direct discharge of RO residuals to a receiving surface water is one possible disposal
alternative. Assume the brine waste stream generated has an arsenic concentration of 265 Mg/L.
Now, assume the state regulatory agency has set an ambient arsenic water quality standard at 20.5
,ug/L for fish consumption. RO will produce approximately 176,000 gallons of brine waste for each
million gallons of treated water. Based upon this scenario, a dilution factor of around 12 is needed
for the receiving water to assimilate the arsenic from the brine stream. In other words, a one million
gallon per day treatment plant would require a dilution flow of 3 cfs to assimilate the 176,000
gallons of waste brine generated. Furthermore, to be discharged directly to a receiving water the
brine stream must pass the WET test. Due to the low arsenic concentration in the RO process
effluent, it is unlikely that the waste brine would fail the WET test (AWWARF, 1998). Therefore,
direct discharge to a receiving water may be a possible disposal alternative for RO treatment
facilities.
The effluent arsenic concentration of 265 //g/L estimated for the RO process is well within
the bounds of the range of TBLLs (50 /ug/L to 1,000 ,ug/L). Therefore, in most cases discharge to
a sanitary sewer will be an acceptable method of disposal for RO residuals containing arsenic. The
appropriate Industrial Protection Program should be checked prior to disposal to verify that the
arsenic concentration of the residuals does not exceed the TBLL.
RO treatment produces high volume liquid waste streams. As a result, land application is
an unlikely candidate for disposal of arsenic containing residuals. The reject water would have to
be applied at a rate matching the evapotransportation requirements of the cover crop grown and is
limited to 41 kg of arsenic per hectare. Therefore, land application is deemed inappropriate for RO
residuals disposal (AWWARF, 1998).
The liquid waste produced by RO treatment can not be sent to a sanitary or hazardous waste
landfill. The high free liquids content makes dewatering uneconomical. The arsenic concentration
is well below TCLP limits which makes hazardous waste landfill disposal unnecessary. Landfill
disposal of RO residuals is inappropriate (AWWARF, 1998).
4-29
-------
4.5 SUMMARY OF RESIDUALS HANDLING AND DISPOSAL OPTIONS
Characteristics of arsenic containing residuals were presented for the following treatment
technologies evaluated within this document: coagulation/filtration; direct filtration; coagulation
assisted microfiltration; enhanced coagulation; lime softening; enhanced lime softening; ion
exchange; activated alumina; reverse osmosis; nanofiltration; ultrafiltration; and microfiltration.
Table 4-1 summarizes the residuals characteristics for each of these treatment processes.
This chapter also evaluated handling and disposal options for water treatment residuals
containing arsenic. Specifically, the following handling options were; presented: gravity thickening;
mechanical dewatering, including filter presses and centrifuges; evaporation ponds and drying beds;
and storage lagoons. Moreover, the following disposal options were evaluated: direct discharge to
surface water; indirect discharge - discharge to sanitary sewer for treatment at POTW; land
application; sanitary landfill disposal; and hazardous waste landfill disposal. Table 4-2 summarizes
the applicability of each of the handling and disposal options presented in this chapter to the
treatment technologies presented throughout this document.
4-30
-------
TABLE 4-1
Summary of Residuals Characteristics1
Treatment Technology
Coagulation/Filtration
Enhanced Coagulation
Direct Filtration
Coagulation Assisted Microfilt ration
Lime Softening
Enhanced Lime Softening
Ion Exchange
Activated Alumina (with Regeneration
and pH optimization)
Activated Alumina (no Regeneration)2
Microfiltration
Ultrafiltration
Nanoflltration
Reverse Osmosis
Waste
Type
Slowdown
Backwash
Slowdown
Backwash
Backwash
Brine
Slowdown
Slowdown
Brine
Brine
Spent
Resin
Reject
Reject
Brine
Brine
Waste
per Million
Gallons Water
Produced
1,400 gallons
10,000 gallons
1,900 gallons
10,000 gallons
20,000 gallons
52,600 gallons
20,000 gallons
20,000 gallons
-
--
-
52,600 gallons
11 1,000 gallons
176,000 gallons
176,000 gallons
Solids
Production per
Million Gallons
Water Produced
1 1 0 pounds
1 0,000 mg/L
160 pounds
10,000 mg/L
45 pounds
45 pounds
103 mg/L
1 ,665 pounds
1.0%
1,665 pounds
1.0%
-
-
445 - 2,450 pounds
-
-
-
-
Arsenic
Concentration
28.0 mg/L
2,930 mg/kg
2,930 mg/kg
21.0 mg/L
2,060 mg/kg
2,060 mg/kg
2.0 mg/L
7,333 mg/kg
744^^
7,333 mg/kg
2.0 mg/L
200 mg/kg
2.0 mg/L
200 mg/kg
10.0 mg/L
14,250 mg/kg
9.5 mg/L
14,250 mg/kg
140 -760 mg/kg
190 Mg/L
315 //g/L
225 jzg/L
260/ug/L
1 Based upon design criteria established in Chapter 3 of this document For general residuals characteristics, see the appropriate section of this
chapter.
2 Dependent upon raw water pH. Lower bound of the solids production is based on a raw water pH of 7, upper bound is based on pH 8. Arsenic
concentration lower bound is based on pH 8, and upper bound is pH 7. pH 7 is closer to the optimal pH (5.5). As a result less resin is needed
to remove the same amount of arsenic. Conversely, because less resin is needed but the same amount of arsenic is removed, the concentration
of the waste stream is increased.
4-31
-------
TABLE 4-2
Summary of Arsenic Residuals Handling and Disposal Options
Treatment Technology
Coagulation/Filtration
Enhanced Coagulation
Direct Filtration
Coagulation Assisted Microfiltration
Lime Softening
Enhanced Lime Softening
Ion Exchange
Activated Alumina (with Regeneration
and pH optimization)
Activated Alumina (no Regeneration)
Microfiltration
Ultrafiltration
Nanofiltration
Reverse Osmosis
Handling Options
GT
•
•
•
•
•
•
•
•
X
•
•
•
•
MD
•
•
•
•
•
•
•
•
X
•
•
•
•
EP
o
0
o
0
o
o
0
o
X
o
o
o
o
SL
o
0
o
0
o
o
o
o
X
o
0
•
•
Disposal Options
DO
X
X
•
X
X
X
X
X
X
o
o
o
0
ID
•
•
•
O
•
•
•
•
X
o
o
o
o
LA
o
o
o
o
0
o
o
o
X
o
•
X
X
so
o
o
o
o
0
o
0
o
o
o
o
X
X
HD
n
a
a
a
a
D
n
a
a
D
D
X
X
GT = gravity thickening, MD = mechanical dewatering, EP = evaporation ponds and drying beds, SL== storage lagoons,
DD = direct discharge, ID = indirect discharge, LA = land application, SD = sanitary landfill disposal, HD - hazardous landfill disposal
• -Yes
o - Yes, with limitations. Disposal may depend upon NPDES limits, TBLLs, TCLP results, or presence of free liquids.
• -Not likely
D - Only as a last resort
x - Not a feasible option
4-32
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4.6 RESIDUALS HANDLING AND DISPOSAL COSTS
The disposal costs presented in this section are based on the handling and disposal options
presented in Sections 4.2 and 4.3, and on the residuals characteristics presented in Section 4.4. Cost
estimates were generated using the Small Water System Byproducts Treatment and Disposal Cost
Document (DPRA, 1993a), and the Water System Byproducts Treatment and Disposal Cost
Document (DPRA, 1993b).
Costs are presented for five water treatment processes: (1) coagulation assisted
microfiltration, (2) ion exchange, (3) activated alumina, (4) reverse osmosis, and (5) greensand
filtration. The disposal options assumed for each of these treatment processes are shown below:
• Coagulation Assisted Microfiltration:
- Mechanical dewatering and non-hazardous landfill disposal
- Non-mechanical dewatering and non-hazardous landfill disposal
• Ion Exchange:
- POTW (indirect) discharge
- Evaporation pond and non-hazardous landfill disposal
- Chemical precipitation and non-hazardous landfill disposal
• Activated Alumina:
- Non-hazardous landfill disposal (systems operating without regeneration)
- POTW (indirect) discharge and non-hazardous landfill disposal
i
• Reverse Osmosis:
- Direct discharge
- POTW (indirect) discharge
- Chemical precipitation and non-hazardous landfill disposal
• Greensand Filtration:
- POTW (indirect) discharge
The handling and disposal options presented represent those selected by EPA during the
development of the arsenic rulemaking decision tree. These disposal options will be used by EPA
during the Regulatory Impact Analysis. Costs for additional disposal options can be estimated using
the DPRA disposal cost documents (DPRA, 1993 a and 1993b). Capital and O&M cost estimates
for the selected handling and disposal options are presented in Figures 4-1 through 4-27. Waste
4-33
-------
volumes upon which the estimates are based are presented in Table 4-1. Capital cost estimates for
non-hazardous landfill disposal are assumed zero. As a result, capital cost estimates are not
provided when landfill is the sole disposal technology (i.e., activated alumina systems choosing not
to regenerate). Capital cost estimates are provided in instances where additional technologies are
employed, for example, mechanical dewatering and non-hazardous landfill disposal. Land costs
are excluded from all capital cost estimates.
4-34
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5.0 POINT-OF-ENTRY/POINT-OF-USE TREATMENT OPTIONS
5.1 INTRODUCTION
Centralized treatment is not always a feasible treatment option, for example, in areas where
each home has a private well or centralized treatment is cost prohibitive. In these instances, point-
of-entry (POE) and point-of-use (POU) treatment options may be acceptable treatment alternatives.
POE and POU systems offer ease of installation, simplify operation and maintenance, and generally
have lower capital costs (Fox, 1989). These systems may also reduce engineering, legal and other
fees typically associated with centralized treatment options. Use of POE and POU systems does not.
reduce the need for a well-maintained water distribution system. In fact, increased monitoring may
be necessary to ensure that the treatment units are operating properly.
Home water treatment can consist of either whole-house or single faucet treatment. Whole-
house, or POE treatment is necessary when exposure to the contaminant by modes other than
consumption is a concern. POU treatment is preferred when treated water is needed only for
drinking and cooking purposes. POU treatment usually involves single-tap treatment.
Section 1412(b)(4)(E) of the 1996 Safe Drinking Water Act (SDWA) Amendments requires
the EPA to issue a list of technologies that achieve compliance with MCLs established under the act.
This list must contain technologies for each NPDWR and for each of the small public water systems
categories listed below:
• Population of more than 50, but less than 500;
• Population of more than 500, but less than 3,300; and
• Population of more than 3,300, but less than 10,000.
The SDWA identifies POE and POU treatment units as potentially affordable technologies,
but stipulates that POE and POU treatment systems "shall be owned, controlled and maintained by
the public water system, or by a person under contract with the public water system to ensure proper
operation and compliance with the maximum contaminant level or treatment technique and equipped
5-1
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with mechanical warnings to ensure that customers are automatically notified of operational
problems."
Research has shown that POE and POU devices can be effective means of removing arsenic
from potable water (Fox and Sorg, 1987; Fox, 1989). Water systems with high influent arsenic
concentrations, i.e., greater than 1 mg/L, may have difficulty meeting MCLs much lower than the
10 to 20 //g/L level. As a result, influent arsenic concentration and othear source water characteristics
must be considered when evaluating POE and POU devices for arsenic removal. To be effective
these devices should work with minimal attention and be relatively inexpensive for the user. Reverse
osmosis, activated alumina, and ion exchange are three treatment techniques that have been
evaluated and shown to be effective. This chapter looks at the removals achieved by each of these
three treatment techniques, and presents total costs for each treatment option.
5.2 VARIABLES AFFECTING REMOVAL EFFICIENCY
5.2.1 Speciation
Arsenic speciation is critical to the removal efficiency of every technology presented in this
document As previously discussed, inorganic arsenic occurs in two primary forms, arsenite (AsIII)
and arsenate (AsV). Arsenite is removed less efficiently because it predominantly occurs in the
uncharged (H3AsO3) state in source waters with a pH of less than 9.0. The dominant arsenate forms
are anionic species, H2AsO42" and HAsO4".
Arsenic removal is dependent upon water chemistry and arsenic speciation. As a result,
identification of the ionic form of arsenic is necessary for selection and design of a removal process.
All technologies discussed in this document remove arsenate more effectively than arsenite.
Therefore, if arsenite is the predominant species present, oxidation to arsenate may be required to
achieve the desired removal.
5.2.2 pH
As previously stated, pH plays a significant role in determining the removal efficiency of a
particular technology. Most processes are relatively unaffected by pH in the range of 6.5 to 9.0.
5-2
-------
However, activated alumina studies have shown the optimum pH for arsenic removal to be between
5.5 and 6.0, and reverse osmosis processes may require pH adjustment to prevent precipitation of
salts on the membrane surface.
Source water pH can be adjusted by addition of dilute acid. However, the treated water also
may need pH adjustment to control corrosion. This can lead to increase in capital, as well as
operations and maintenance costs. In POE and POU systems there is an added concern of monitoring
to insure that pH levels in treated water are safe.
5.2.3 Co-occurrence
Co-occurrence of inorganic contaminants, such as sulfate and silica, as well as suspended
solids, can cause interference with arsenic removal. Sulfate is preferentially adsorbed relative to
arsenic by ion exchange processes. This preference can result in another phenomenon known as
peaking, which occurs when arsenic is displaced on the resins by the sulfate causing effluent
concentrations in excess of the influent levels.
A slight decrease in activated alumina performance has been seen in waters with high sulfate
concentration, however, the effect is not as great as in ion exchange processes. At higher treatment
pH levels silica may also be preferred relative to arsenic.
5.3 POE/POU DEVICE CASE STUDIES
Several field studies conducted to evaluate the effectiveness of POE and POU treatment units
for arsenic removal indicate that POE and POU systems can be effective alternatives to centralized
treatment options. These studies evaluated reverse osmosis (RO), activated alumina (AA), and ion
exchange (IX) processes.
The following sections present the results of two of these studies. Table 5-1 summarizes
source water quality and influent arsenic concentrations. Table 5-2 summarizes the arsenic removals
achieved by each of the technologies evaluated.
5-3
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TABLE 5-1
Source Water Summary - Point-of-Use Case Studies
Source Water Characteristic
Alkalinity
Arsenic
Calcium
Chloride
Fluoride
Iron
Magnesium
Manganese
PH
Silica
Sodium
Sulfate
Total Dissolved Solids
Total Hardness
Turbidity
Concentration1
56 - 206
<0.005- 1.16
8.9 - 22
<10
0.6 - 5.2
<0. 1-2.5
5.3 - 10.6
<0.6
7.4 - 8.3
NA
4.4 - 62
<15
<1,500
109 - 547
0.24 - 0.48
1 All concentrations are given in mg/L, except turbidity (NTU) and pH units. Note measurements for
each parameter were not taken at each test site.
TABLE 5-2
Observed Arsenic Removal by Technology for POE and POU Units
Treatment Option
Reverse Osmosis
(low pressure)
Reverse Osmosis
(high pressure)
Activated Alumina
Ion Exchange
Arsenic
Removal
(%)
50-86
50-80
NA
NA
Quantity of Water
Treated When MCL
Exceeded* (gal)
50-332
684
0-15,427
0-16,254
Time On-Line Until
MCL Exceeded*
(days)
35-225
350
0-1,226
0-1,471
' Based upon EPA DWRD study conducted in Fairbanks, Alaska and Eugene, Oregon only,
5-4
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5.3.1 Case Study 1: Fairbanks, Alaska and Eugene, Oregon
The EPA Drinking Water Research Division (DWRD) conducted field POU studies in
Fairbanks, Alaska and Eugene, Oregon (Fox and Sorg, 1987; Fox, 1989). Pilot systems were
installed in two homes in each community, and each system consisted of an activated alumina bed,
ion exchange bed and reverse osmosis system. Influent arsenic concentrations ranged from 0.05 to
1.16 mg/L and was believed to be naturally occurring.
The RO systems at each of the four locations performed well upon start-up, achieving 60 to
80% removal of arsenic. Over time, however, the arsenic removal efficiency decreased to 50% or
less. At the initial removal efficiencies, effluent arsenic levels met the current MCL of 50 ppb, but
over time failed to sufficiently reduce the levels to below the MCL.
Low-pressure RO units (40 - 60 psig) consistently achieved greater than 50% removal, but
with high influent arsenic concentrations much higher efficiency is necessary to achieve the MCL.
The high-pressure units (196 psig) operated for 350 days and produced 684 gallons of treated water
which met the MCL.
The IX beds evaluated were 1 cubic foot in size and were filled with a strong base anion
exchange resin (Dowex-SBR). The DC beds effectively reduced arsenic levels to below the MCL,
but required pre-treatment to ensure effective removal. This involved regeneration and chemical
treatment of the resin to the chlorine form.
The AA beds were identical to the DC beds with the exception that they were filled with
granular activated alumina (Alcoa-Fl). The AA beds effectively reduced arsenic levels, however
they required pre-treatment to reduce the pH to 5.5 - 6.0. Regeneration involved passing a sodium
hydroxide solution through the tank, rinsing with clean water, and then treating with dilute sulfuric
acid. Improperly treated alumina performed poorly initially (30 to 40 percent removal), and
performance significantly deteriorated over time (5 to 20 percent removal). Proper pre-treatment,
however, allowed for efficient operation periods of nearly one year or greater.
5.3.2 Case Study 2: San Ysidro, New Mexico
A field study was also conducted in San Ysidro, NM to evaluate the effectiveness of POU
RO units. This work is documented in several sources (Thomson and O'Grady, 1998; Fox, 1989;
5-5
-------
Fox and Sorg, 1987; and Clifford and Lin, 1985). San Ysidro source water is from an infiltration
gallery under the local river banks and contains 5.2 mg/L fluoride and 0.23 mg/L arsenic. The water
is also high in other inorganic contaminants, including iron (2.5 mg/L), manganese (0.6 mg/L) and
total dissolved solids (1,500 mg/L). San Ysidro is a small community with limited financial
resources, and central treatment was not a viable treatment option. San Ysidro applied to the DWRD
for a cooperative agreement to evaluate POU RO treatment for the entire village. The project was
funded in August 1995, and seventy-three units were initially purchased and installed in homes,
restaurants, gas stations and municipal buildings.
Arsenic concentrations in the source water were consistently reduced from 0.068 mg/L to
0.02 mg/L to less than the detectable limit (0.005 mg/L). Other contaminants were'also effectively
removed, including manganese (80 percent), iron (85 percent), and TDS (95 percent). Based upon
the manufacturer's literature, it appeared that the units were operating at an approximate recovery
rate of 25%, i.e., for every 100 gallons of influent, 25 gallons of treated water are produced.
The water supply for San Ysidro is chlorinated at the wellheaid. As a result, a carbon pre-
filter was installed with each unit to remove residual chlorine and particulates to prevent membrane
fouling. A carbon post-filter was installed for polishing treated water. Since the conclusion of the
study, the village has assumed ownership of the units and is now responsible for their maintenance.
5.4 REVERSE OSMOSIS
Reverse osmosis (RO) is a separation process that utilizes a membrane system to reject
compounds based upon molecular properties. Water molecules pass through the membrane, but
most contaminants, including arsenic, are rejected by the membranes. While a portion of the feed
water passes through the membrane, the rest is discharged with the rejected contaminants in a
concentrated stream. Membrane performance can be adversely affected by the presence of turbidity,
iron, manganese, scale-producing compounds, and other contaminants. A detailed discussion of the
RO removal mechanism is presented in Chapter 2.
POU RO systems can be operated at both high (approximately 200 psig) and low (40 - 60
psig) pressures. High pressure RO devices typically operate at a product-to-reject water ratio of 1
5-6
-------
to 3 (Fox, 1989), and require a booster pump to achieve the desired operating pressure. Low
pressure RO devices are less efficient and operate with a product-to-reject water ratio of about 1 to
10 (Fox, 1989). This can be a significant deterrent to RO treatment in dry regions or regions with
frequent water shortages.
Manufacturer and laboratory data suggest greater than 95% removal of arsenate by RO
systems, and slightly less (75%) removal of arsenite. Field studies indicate that greater than 50%
removal is possible, but data are inconclusive much beyond those levels. Accordingly, water
systems with high influent arsenic concentrations, i.e., greater than 1.0 mg/L, may consider other
POE and POU treatment options.
5.4.1 Cost Estimates
The EPA document, Cost Evaluation of Small System Compliance Options - Point-ofUse
and Point-of-Entry Treatment Units (Cadmus Group, 1998), was used to estimate POE and POU RO
treatment costs. Costs are presented in Figure 5-1, and are based upon the following assumptions:
Average household - 3 individuals, 1 gallon each per day, 1,095 gallons per year.
Annual treatment -1,095 gallons (POU), 109,500 gallons (POE).
Minimally skilled labor - $14.50 per hour (population less than 3,300 individuals).
Skilled labor - $28.00 per hour (population greater than 3,300 individuals).
Life of unit - 5 years (POU), 10 years (POE).
Duration of cost study -10 years (therefore, two POU devices per household).
Cost of water meter and automatic shut-off valve included.
No shipping and handling costs required.
Volume discount schedule - retail for single unit, 10% discount for 10 or more units,
15% discount on more than 100 units.
Installation time -1 hour unskilled labor (POU), 3 hours, skilled labor (POE).
O&M costs include maintenance, replacement of pre-filters and membrane
cartridges, laboratory sampling and analysis, and administrative costs.
5-7
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5.5 ION EXCHANGE
Ion exchange (IX) has been used effectively for the removal of arsenic. There are two types
of DC systems, anionic and cationic. IX for arsenic removal is typically anionic since arsenic present
in natural waters is generally anionic. IX is a process wherein an ion in the solid phase (e.g., a
synthetic resin) is exchanged for an ion in the source water. To accomplish this exchange, the source
water is passed through the IX bed in either a downflow or upflow mode until the resin is exhausted.
Exhaustion occurs when unacceptable levels of the contaminant are observed in the bed effluent.
In POE and POU systems water often sits in the beds for extended periods of time because the
required flow is not constant. It is possible that this leads to better than average removal of arsenic.
Further research is needed to explain this effect. The IX removal mechanism is discussed in greater
detail in Chapter 2.
5.5.1 Cost Estimates
The Cadmus Group (1998) document was again used to estimate POE and POU treatment
costs. Costs are presented in Figure 5-2, and are based upon the following assumptions:
Average household - 3 individuals, 1 gallon each per day, 1,095 gallons per year.
Annual treatment -1,095 gallons (POU), 109,500 gallons (POE).
Minimally skilled labor - $14.50 per hour (population less than 3,300 individuals).
Skilled labor - $28.00 per hour (population greater than 3,300 individuals).
Life of unit - 5 years (POU), 10 years (POE).
Duration of cost study -10 years (therefore, two POU devices per household).
Cost of water meter and automatic shut-off valve included.
No shipping and handling costs required.
Volume discount schedule - retail for single unit, 10% discount for 10 or more units,
vendor retains 30% profit on more than 100 units.
Installation time -1 hour unskilled labor (POU), 3 hours, skilled labor (POE);
O&M costs include maintenance, replacement of pre-filters and resin cartridges,
laboratory sampling and analysis, and administrative costs.
5-9
-------
5-10
-------
5.6 ACTIVATED ALUMINA
Activated alumina (AA) can be used in packed beds to remove inorganic contaminants,
including arsenic, from source water. In the AA process, contaminants are exchanged with the
hydroxide ions on the alumina surface. Depending upon pH, AA can act as either an anionic (pH
greater than 8.2) or cationic (pH less than 8.2) exchange process, but is rarely used as a cationic
process for water treatment.
Arsenic removal by AA has been shown to be most effective near pH of 5.5 to 6.0. Most
water systems will need some type of pH adjustment to accommodate this requirement. For POE
and POU systems this can be accomplished by treating the AA bed with dilute sulfuric acid. This
minimizes the possibility of unsafe treated water due to acidic pH, as well as the likelihood that
additional pH adjustment would be necessary to raise the pH after treatment.
5.6.1 Cost Estimates
The Cadmus Group (1998) document was again used to estimate POE and POU treatment
costs. Costs are presented in Figure 5-3, and are based upon the following assumptions:
Average household - 3 individuals, 1 gallon each per day, 1,095 gallons per year.
Annual treatment -1,095 gallons (POU), 109,500 gallons (POE).
Minimally skilled labor - $14.50 per hour (population less than 3,300 individuals).
Skilled labor - $28.00 per hour (population greater than 3,300 individuals).
Life of unit - 5 years (POU), 10 years (POE).
Duration of cost study -10 years (therefore, two POU devices per household).
Cost of water meter and automatic shut-off valve included.
No shipping and handling costs required.
Volume discount schedule - retail for single unit, 10% discount for 10 or more units,
vendor retains 30% profit on more than 100 units.
Installation time -1 hour unskilled labor (POU), 3 hours, skilled labor (POE).
O&M costs include maintenance, replacement of pre-filters and resin cartridges,
laboratory sampling and analysis, and administrative costs.
5-11
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6.0 REGIONALIZATION
6.1 BACKGROUND
The term regionalization is used to define the process of purchasing and transporting water
from one community to another. In effect, regionalization expands the region served by a water
distribution system. Numerous economic, geographic, and operational factors can influence the
decision to implement regionalization, including: (1) the availability of water; (2) water quality; (3)
geography; and (4) economic factors.
Thriving communities that rapidly expand can easily outgrow their water source and find
themselves faced with water shortage problems. To alleviate this problem, communities may decide
to purchase water from other available sources in the region or neighboring communities. Water
quality also plays a role in the decision making process. If a community's source water is
contaminated, it may be cheaper for the community to purchase water from another community
rather that treat its own water source. In some cases, contaminated water cannot be sufficiently
treated and a community may be faced with a choice to establish a new water source or to purchase
water from a neighboring community.
Economic factors affecting the decision for regionalization include design, materials,
construction, land, labor, and operational costs. Design costs include the engineering fees paid for
the design of the regionalization system. Material costs include piping, fittings, gaskets, bends,
valves, booster stations, pumps, and cathodic protection, among others. Construction costs include
the costs associated with equipment rental and operation, excavation, backfilling, compaction, and
landscaping. Land costs include the land required for the placement of the piping and booster
stations, and the land required for the pipeline right-of-way. Labor costs include equipment
operators, laborers, superintendents, and site engineer. Operational costs include energy costs,
replacement parts, calibration, retrofitting, and operator costs.
The geographic location of a community will greatly affect the economic feasibility of
regionalization. The distance from the water source will affect construction and equipment costs,
and hilly or mountainous terrain can add significant design and construction costs. In addition,
6-1
-------
obtaining right-of-way for pipelines and booster stations may be a significant factor in the decision
making process.
Additional factors include the lack of available water sources or change in the source
availability due to increased drawdown of groundwater, droughts affecting reservoirs, and other
man-made or natural changes to the water source. Increased per capita' water use can increase the
demand for a larger water source or a new one, which also affects the decision process when
considering regionalization. Political issues associated with natural drainage boundaries, the desire
to avoid dependence on a single water source, and the reliability of the water source supply can also
affect the decision for regionalization.
6.2 COST ESTIMATES
Estimated costs for regionalization of drinking water were developed based upon three
construction cost estimating sources: (1) the 1995 Environmental Cost Handling Options and
Solutions (ECHOS); (2) the 1994 National Construction Estimator (NCE); and (3) the 1997 ^Means
- Site Work and Landscape Cost Data, 16th edition. The data collected for the cost estimates are
included in Appendix H. Table 6-1 compares the cost of regionalization using reinforced concrete
pipe, HDPE pipe, and ductile iron pipe as the conveyance conduit material. The following
assumptions were made for the purpose of estimating regionalization costs:
• A 92" wide by 120" deep trench was excavated for the placement of the conveyance
conduit. The width of the trench allows hand compaction around the pipe, the depth is
an average depth.
• Type of soil was not taken into consideration, and no-rock excavation was assumed.
• 12" of fine gravel and sand were used to underlay the pipe.
• 3-48 magnesium anodes, at a spacing of 5 per mile, are assumed for the ductile iron pipe
cathodic protection.
• The costs developed do not include costs associated with fittings, bends, gaskets, tees,
etc. These costs may vary greatly depending upon the topography of the site.
6-2
-------
Air valves were assumed at 2 valves per mile. The location of air valves is also
dependent on the topography of the site; valves are usually located at the high points.
One booster station with a 100 GPM, 150' Head, and a 10 HP centrifugal pump, is
assumed every two miles along the pipeline. The spacing and size of booster stations are
site dependent.
The cost of land purchase is not included.
No escalation factors were used.
The cost estimates do not include design costs, contractor profit and/or additional costs.
Table 6-1
Regionalization Cost Estimates
Pipe Construction
Reinforced
Concrete - Class 3
HDPE
Ductile Iron
Pipe Diameter
(inches)
12
15
24
12
16
24
12
16
24
1995 ECHOS
(S/mile)
$331,399
$341,906
$402,098
NA
NA
NA
$409,706
$456,170
$664,202
1994 NCE
(S/mile)
$206,707
$236,117
$315,950
$290,659
$376,195
$597,955
$389,188
$497,639
$698,543
1997 '"MEANS
(S/mile)
$286,128
$301,176
$374,568
$329,688
$356,088
$501,288
$428,956
$518,716
$685,036
NA = Not Available
6-3
-------
-------
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7-9
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APPENDIX A
VERY SMALL SYSTEMS
CAPITAL COST BREAKDOWN SUMMARIES
-------
-------
Table Al - VSS Document Capital Cost Breakdown for Membrane Processes
Component
Equipment
Installation
Sitework/Interface Piping
Standby Power
OH&P
Legal & Admin
Engineering
Contingencies
Total
Capital Cost
Factor
1.0000
0.2500
0.0750
0.0625
0.1665
0.0416
0.1596
0.0000
1.7552
Percent of Total
Capital Cost
56.97%
14.24%
4.27%
3.56%
9.49%
2.37%
9.09%
0.00%
100.00%
Capital Cost
Breakdown
Category
P
c
c
c
e
e
e
c
Table A2 - VSS Document Capital Cost Breakdown for Ion Exchange Processes
Component
Equipment
Installation
Sitework/Interface Piping
Standby Power
OH&P
Legal & Admin
Engineering
Contingencies
Total
Capital Cost
Factor
1.0000
0.3000
0.0780
0.0650
0.1732
0.0433
0.1659
0.0000
1.8254
Percent of Total
Capital Cost
54.78%
16.43%
4.27%
3.56%
9.49%
2.37%
9.09%
0.00%
100.00%
Capital Cost
Breakdown
Category
P
c
c
c
e
e
e
c
Table A3 - VSS Document Capital Cost Breakdown for Chlorination
Component
Equipment
Installation
Sitework/Interface Piping
Standby Power
OH&P
Legal & Admin
Engineering
Contingencies
Total
Capital Cost
Factor
1.0000
0.1500
0.0690
0.0575
0.1532
0.0383
0.1468
0.0000
1.6148
Percent of Total
Capital Cost
61.93%
9.29%
4.27%
3.56%
9.49%
2.37%
9.09%
0.00%
100.00%
Capital Cost
Breakdown
Category
P
c
c
c
e
e
e
c
International Consultants, Inc.
Contract 68-C6-0039
A-1
November 1999
VSS Document Capital Cost Breakdown
-------
Table A4 - VSS Document Capital Cost Breakdown for Potassium Permanganate Feed
Component
Equipment
Installation
Sitework/Interface Piping
Standby Power
OH&P
Legal & Admin
Engineering
Contingencies
Total
Capital Cost
Factor
1.0000
0.1000
0.0660
0.0550
0.1465
0.0366
0.1404
0.0000
1.5446
Percent of Total
Capital Cost
64.74%
6.47%
4.27%
3.56%
9.49%
2.37%
9.09%
0.00%
100.00%
Capital Cost
Breakdown
Category
P
c
c
c
e
e
e
c
Table AS - Typical VSS Document Capital Cost Breakdown
Component
Equipment
Installation
Sitework/Interface Piping
Standby Power
OH&P
Legal & Admin
Engineering
Contingencies
Total
Capital Cost
Factor
1.0000
0.3000
0.0780
0.0650
0.1732
0.0433
0.1659
0.0000
1.8254
Percent of Total
Capital Cost
54.78%
16.43%
4.27%
3.56%
[_ 9.49%
2.37%
9.09%
0.00%
100.00%
Capital Cost
Breakdown
Category
P
c
c
c
e
e
e
c
International Consultants, Inc.
Contract 68-C6-0039
A-2
November 1999
VSS Document Capital Cost Breakdown
-------
APPENDIX B
WATER MODEL
CAPITAL COST BREAKDOWN SUMMARIES
-------
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-------
-------
APPENDIX D
COST EQUATIONS AND CURVE FITS
FOR REMOVAL AND ACCESSORY COSTS
-------
-------
1.E+06
1.E+05 -
o
O
'a.
ro
O
1.E+04|,
1.E+03 J—
0.01
Figure D-1
Capital Costs for Preoxidation with 1.5 mg/L Permanganate
0.1
Equation
1. y = 9711
2. y = 61277x°174
R2 = 0.9974
a. y = 9085x + 626
1 10
Design Flow (mgd), x
100
Applicable Flow Range
1. 0.01-1 mgd
2. 10-430 mgd
a. 1-10 mgd
1000
D-1
-------
1.E+07
1.E+06 -
•K 1.E+05 -
oS
o
Figure D-2
O&M Costs for Preoxidation with 1.5 mg/L Permanganate
1.E+04 -
A A
1.E+03
0.001
0.01
Equation
0.1 1 10
Average Rows (mgd), x
100
1. y = -5608X2 + 4182.2x + 7237.4
R2 = 0.9875
2. y = -0.1016X2 + 6208.7x + 13443
R2 = 1
a. y = 8059x + 5115
Applicable Flow Range
1. 0.003 - 0.03 mgd
2. 4.5-270 mgd
a. 0.03 - 4.5 mgd
1000
D-2
-------
1.E+06
1.E+05 •
Figure D-3
Capital Costs for Preoxidation with 3.0 mg/L Permanganate
tn
~
-------
1.E+07
1.E+06 -
1.E+05 -
o
O
1.E+04 -
1.E+03
0.001
Figure D-4
O&M Costs for Preoxidation with 3.0 mg/L. Permanganate
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Equation
1. y = 22244X2 + 2826.x + 7221
R2 = 0.9937
a. y = 14166x + 6022
Applicable Flow Range
1. 0.003 - 0.03 mgd
2. y = -0.128X2 + 12379x + 14064 2. 4.5 - 270 mgd
R2 = 1
a. 0.03 - 4.5 mgd
D-4
-------
1.E+06
1.E+05
Figure D-5
Capital Costs for Preoxidation with 5.0 mg/L Permanganate
"25
8
-i
"a.
CO
O
1.E+04,L.
-M-
1.E+03
0.01
0.1
1 10
Design Flows (mgd), x
100
1000
Equation
1. y = 9711
2. y = 77113x01691
R2 = 0.9991
a. y=11568x-1857
Applicable Flow Range
1. 0.01-1 mgd
2. 10-430 mgd
a. 1-10 mgd
D-5
-------
1.E+07
1.E+06 •
o
O
1.E+04 -
1.E+03
0.001
Figure D-6
O&M Costs for Preoxidation with 5.0 mg/L Permanganate
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Equation
1. y = 21186X2 + 12736X + 7040
R2 = 0.995
Applicable Flow Range
1. 0.003 - 0.03 mgd
2. y = -0.141 Ix2 + 20608X + 14723 Z 4'5 ' 27° mgd
R2 = 1
a. y = 22484x^6277 a. 0.03-4.5 mgd
D-6
-------
1.E+06
1.E+05 •
o
O
3s
'a.
to
O
1.E+04 •
1.E+03
0.01
Figure D-7
Capital Costs for Preoxidation with 1.5 mg/L Chlorine
0.1
1 10
Design Flows (mgd), x
100
Equation Without Housing
1. y = 7277
2. y = 15290x°'6554
R2 = 0.9892
a. y = 6875x + 402
Equation With Housing
3. y = 14560
4. y = 18538xa6554
R2= 0.9892
b. y = 7698x + 6862
Applicable Flow Range
2. 10-430 mgd
a. 1-10 mgd
Applicable Flow Range
3. 0.01 -1 mgd
4. 10-430 mgd
b. 1 -10 mgd
1000
D-7
-------
1.E+06
1.E+05 -
-------
1.E+07
1.E+06 -
CO
O 1.E+05 -I
".
CO
O
o
1.E+04 -
1.E+03
0.01
Figure D-9
Capital Costs for Preoxidation with 3.0 mg/L Chlorine
o o oo
0.1
1 10
Design Flows (mgd), x
100
Equation Without Housing
1. y = 7277
2. y = 28225x0597
R2 = 0.9889
a. y = 11591x-4314
Equation With Housing
3. y = 14560
4. y = 38026x0'5762
R2 = 0.9923
b. y = 14306x + 254
Applicable Flow Range
1. 0.01-1 mgd
2. 10-430 mgd
a. 1-10 mgd
Applicable Row Range
3. 0.01 -1 mgd
4. 10-430 mgd
b. 1-10 mgd
1000
D-9
-------
1.E+06
1.E+05 -
(3
1.E+04 -
1.E+03
0.001
Figure D-10
O&M Costs for Preoxidation with 3.0 mg/L Chlorine
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Equation
1. y = 8479.4x + 1161.8
R2 = 1
Applicable Flow Range
1. 0.003 - 0.03 mgd
2. y = -0.4237X2 + 1468.8x + 25798 2. 4.5 - 270 mgd
R2 = 0.9987
a. y = 6808x + 1764
a. 0.03-4.5
D-10
-------
1.E+07
Figure D-11
Capital Costs for Preoxidation with 5.0 mg/L Chlorine
0.1
1 10
Design Flows (mgd), x
100
Equation Without Housing
1. y = 7277
2. y = 48678X0'5496
R2 = 0.9892
a. y = 17655x-10378
Equation With Housing
3. y = 14560
4. y = 67463x°'5157
R2 = 0.9892
b. y = 22959x - 8399
D-ll
Applicable Flow Range
1. 0.01 - 1 mgd
2. 10-430 mgd
a. 1-10 mgd
Applicable Flow Range
3. 0.01 -1 mgd
4. 10-430 mgd
b. 1-10 mgd
1000
-------
1.E+06
1.E+05 •
"w
o
O
00
O
1.E+04
1.E+03
0.001
Figure D-12
O&M Costs for Preoxidation with 5.0 rng/L Chlorine
0.01
0.1 ' 1 10
Average Flow (mgd), x
100
Equation
1. y = 14132x +1161.8
R2 = 1
2. y=10255xa7241
R2 = 0.9968
a. y = 5851x + 4143
Applicable Flow Range
1. 0.003 - 0.03 mgd
2. 4.5-270 mgd
a. 0.03-4.5 mgd
1000
D-12
-------
Figure D-13
Capital Costs for Coagulation/Filtration
1e+9
1e+4
0.01
1 10
Design Flows (mgd), x
1000
Accessories Costs
a. y = 821328x-19706
b. y = 331382x +53484
1. y = 23072 + 452772x - 592228X2
R2 = .99
2. y = 117680 + 329248x - 62062X2
R2 = .99
3. y = 1967171 + 140479X - 46.61X2
R2 = .99
Removal Costs
c. y = 796179x +105382
d. y = 1187734x- 506079
4. y = 68372 + 1341847x - 1755639X2
R2 = .99
5. y = 152413 + 656296X - 127054X2
R2 = .99
6. y = 6973081+ 440928X - 111x2
R2 = .99
Applicable Flow Range
0.1 -0.27 mgd
1 -10mgd
0.01 - 0.1 mgd
0.27-1.0 mgd
10-430 mgd
0.1 -0.27 mgd
1 -10 mgd
0.01 - 0.1 mgd
0.27-1.0 mgd
10-430 mgd
D-13
-------
Figure D-14
O&M Costs for Coagulation/Filtration
1e+8
V)
o
O
1e+7 -
1e+6 -.
1e+5 ,
1e+4 .
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 25612X + 2502
b. y = 21464x+1787
1. y = 2515 + 35186X - 349122X2
R2 = 99
2. y = 3144 + 18273x - 1610X2
R2 = .99
3. y = 31465 + 14869X + 0.0082X2
R2 = .99
Removal Costs
c.y = 241727x +11633
d.y = 77729x + 21831
4. y = 14772 + 206708X - 2050991X2
R2 = 0.99
5. y = 26579+ 69014X - 12422X2
R2 = 0.99
6. y = 128650+ 54006x - 3.14X2
R2 = 0.99
Applicable Flow Range
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.08S - 0.36 mgd
4.5 - 270 mgd
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
D-14
-------
Figure D-15
Capital Costs for Coagulation Direct Filtration
1e+8 q
1e+4
0.01
0.1 1 10 100
Design Flows (mgd), x
1000
Accessories Costs
a. y = 900104x-40976
b. y = 331382x +53484
1. y = 16427 + 536267X - 2101926X2
R2 = .99
2. y = 117680 + 329248X - 62062x2
R2 = .99
3. y = 1967171 + 140479X - 46.61 x2
R2 = .99
Removal Costs
c.y = 1029729x + 42342
d. y = 758882X - 77053
4. y = 48683 + 1589204X - 6228853X2
R2 = .99
5. y = 152609 + 655401X - 126181X2
R2 = .99
6. y= 5721101+ 180077X-101X2
R2 = .99
Applicable Flow Range
0.1 - 0.27 mgd
1 -10 mgd
0.01 - 0.1 mgd
0.27-1.0 mgd
10-430 mgd
0.1 - 0.27 mgd
1 -10 mgd
0.01 - 0.1 mgd
0.27-1.0 mgd
10-430 mgd
D-15
-------
Figure D-16
O&M Costs for Coagulation Direct Filtration
1e+8 ,
1e+7 ,
1e+6 ,
1e+5 -.
1e+4 -
1e+3 ,
0.001
0.01 0.1 1 10 100 1000
Average Flows (mgd), x
Accessories Costs
a. y = 57798X - 266.6
b. y = 21464x + 1787
1. y = 1219 + 9642x - 52498x2
R2 = .99
2. y = 3144 + 18273X - 1610X2
R2 = .99
3. y = 31465 + 14869x + 0.0082X2
R2 = .99
Removal Costs
c. y = 432745X - 4794
d.y = 58094x + 28900
4. y = 7161 + 56646X - 30841 Ox2
R2 = 0.99
5. y = 26579+ 69014x - 12422X2
R2 = 0.99
6.y= 117979 +38311x-2.77x2
R2 = 0.99
Applicable Flow Range
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4,5 - 270 mgd
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
D-16
-------
Figure D-17
Capital Costs for Coagulation Assisted Microfiltration
1e+9
1e+4
0.01
1 10
Design Flows (mgd), x
100
1000
Equation
a.y = 1611656x + 17173
b. y = 520505.1 x +299748
1.y= 680212.8X0'5814
R2 = 0.99
2. y = 820252.8x° 4546
R2 = 0.99
3.y = 2129119 + 921471x + 320X2
R2 = 0.99
Applicable Flow Range
1.0.01-0.1 mgd
2. 0.27-1 mgd
3. 10-430 mgd
4. 0.01 - 0.1 mgd
• Removal Cost
o Accessories Cost
c. y = 2343200X + 228653
d.y = 103081 Ox+ 1067733
4. y = 94324 + 4880036x - 11935465X2
R2 = 0.99
5. y = 273143 + 2308991x - 483591X2
R2 = 0.99
6. y = 2712738 + 279433.6x - 22.75x2
R2 = 0.99
5. 0.27 -1 mgd
6. 10-430 mgd
a. 0.1 -0.27 mgd
b. 1 -10 mgd
D-17C. 0.1 -0.27 mgd
d. 1 -10 mgd
-------
CO
O
u
§
1e+8
1e+7 .
1e+6 -
1e+5 -
1e+4 -
1e+3
0.001
Figure D-18
O&M Costs for Coagulation Assisted Microfiltration
0.01
0.1 1 ' 10
Average Flows (mgd), x
100
1000
Equation
a.y = 22355 + 216800x
b.y= 55196+81155X
1. y = 20308 + 285019x
R2 = 0.99
2.y= 27685+ 157575X
R2 = 0.99
3.y = 46245 + 83144x
R2 = 0.99
Applicable Flow Range
1.0.003-0.03 mgd
2. 0.09 - 0.36 mgd
3. 4.5 - 270 mgd
4. 0.003 - 0.03 mgd
• Removal Cost
o Accessories Cost
c. y = 2662 + 124224.6X
d.y = 16166 + 34945.8x
4.y = 2612 + 125905.1x
R2 = 0.99
5.y= 8875 + 55198.1X
R2 = 0.99
6.y= 61914 + 24779.5X
R2 = 0.99
5. 0.09 - 0.36 mgd
6. 4.5 - 270 mgd
a. 0.03 - 0.09 mgd
b. 0.36 - 4.5 mgd
c. 0.03 - 0.09 mgd
- 4.5 mgd
-------
Figure D-19
Capital Costs for Enhanced Coagulation/Filtration
1e+9
1e+8 ,
1e+7 -.
CO
"w ie+6
o
O
& 1e+5
O
1e+4 ,
1e+3
0.01
0.1
1 10
Design Flows (mgd), x
100
1000
Accessories Costs
y = 0
Removal Costs
a. y = 89662x-71592
1. y = 7027 + 14922x - 3879x2
R2 = .99
2. y = 756410 + 6937x - 7.55x2
R2 = .99
Applicable Flow Range
0.01-430 mgd
1 -10mgd
0.01 -1 mgd
10-430 mgd
D-19
-------
1.E+07
1.E+06 -
-c 1.E+05
i
"w
o
O
o3 1.E+04
O
1.E+03 -
1.E+02
0.001
Figure D-20
O&M Costs for Enhanced Coagulation/Filtration
0.01
0.1 1 10
Average Flow (mgd), x
100
Accessories Costs
y = 0
Removal Costs
a. y = 17479x-3483
1. y = -302.31x2
R2=1
2. y = -0.3972X2 + 14950x + 7906
R2=1
Applicable Flow Range
0.003 - 270 mgd
0.36-4.5 mgd
0.003 - 0.36 mgd
4.5 - 270 mgd
1000
D-20
-------
Figure D-21
Capital Costs for Lime Softening
1e+9
1e+5
0.01
1 10
Design Flows (mgd), x
100
1000
Accessories Costs
a. y = 331477X+52533
1. y = 104084 + 380454X - 100528X2
R2 = .99
2. y = 1967171 + 140479X-46.61X2
R2 = .99
Removal Costs
b. y = 1199712x- 519736
3. y = 204540 + 421712x + 53724X2
R2 = 0.99
4. y = 8187348 + 330434x - 143X2
R2 = 0.99
Applicable Flow Range
1 -10 mgd
0.01 -1 mgd
10-430 mgd
1-10 mgd
0.01 -1 mgd
10-430 mgd
D-21
-------
Figure D-22
O&M Costs for Lime Softening
1e+8
co
«*-•
tn
O
O
1e+7 -
1e+6 ,
1e+5 ,
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 21470x+1762
1. y = 2931 + 20690X - 6849X2
R2 = .99
2. y = 31465 + 14869x + 0.0082X2
R2 = .99
Removal Costs
b.y = 101406x + 68702
3. y = 62903 + 128682x - 38736x2
R2 = 0.99
4.y = 161285 + 80863x- 6.96x2
R2 = 0.99
Applicable Flow Range
0.36 - 4.5 mgd
0.003 - 0.36 mgd
4.5 - 270 mgd
0.36 - 4.5 mgd
0.003 - 0.36 mgd
4.5 - 270 mgd
D-22
-------
Figure D-23
Capital Costs for Enhanced Lime Softening
1e+8 -,
1e+7 -
1e+6 ,
"
O
O
3 1e+5 4
'a.
CD
O
1e+4 -
1e+3
0.01
0.1
1 10
Design Flows (mgd), x
100
1000
Accessories Costs
y = 0
Removal Costs
a. y = 133487x-100318
1. y = 7727 + 48288X - 22847X2
R2 = 0.99
2. y = 969164 + 26611x - 7.5X2
R2 = 0.99
Applicable Flow Range
0.01-430 mgd
1 -10 mgd
0.01 -1 mgd
10-430 mgd
D-23
-------
1.E+07
1.E+06 -
•C? 1.E+05 -
S| 1.E+04 -
1.E+03 -
Figure D-24
O&M Costs for Enhanced Lime Softening
1.E+02
0.001
0.01
0.1 1 10
Average Flow (mgd), x
100
1000
Accessories Costs
y = 0
Removal Costs
a. = 29145x + 1094
. y = 2144.7xz +
R2 = 0.9999
2. y = -0.4528X2 + 26067x + 14956
R2 = 1
Applicable Flow Range
0.003 - 270 mgd
0.36 - 4.5 mgd
0.003 - 0.36 mgd
4.5 - 270 mgd
D-24
-------
Figure D-25
Capital Costs for Activated Alumina
1e+9
1e+4
0.01
0.1 1 10
Design Flows (mgd), x
100
1000
Accessories Costs
a. y = 2x106x-178252
b.y = 128996x+554580
1. y = 18798 + 94009x +1862274X2
R2 = .99
2. y = 341549 + 319454x + 22573X2
R2 = .99
3. y = 821815 + 102117x + 15.61X2
R2 = .99
Removal Costs
c. y = 7296957X - 620445
d.y = 278725x + 2183518
4. y = 43863 + 219335X + 4345438X2
R2 = 0.99
5. y = 985627 + 1301173X+ 175442.8X2
R2 = 0.99
6. y = 977094 + 399034x + 33.3X2
R2 = 0.99
D-25
Applicable Flow Range
0.1 - 0.27 mgd
1 -10 mgd
0.01-0.1 mgd
0.27-1.0 mgd
10-430 mgd
0.1 - 0.27 mgd
1 -10 mgd
0.01 -0.1 mgd
0.27-1.0 mgd
10-430 mgd
-------
Figure D-26
Capita! Costs for Activated Alumina with No Regeneration
1e+7
1e+4
0.01
Design Flows (mgd), x
Cost Equation
1.y = 11862+1744516x
R2 = 0 .99
2. y = 8490+1013936x
R2 = 0.99
Cost Assumptions
1. 2 columns with infrastructure
2.1 column with infrastruture
D-26
-------
Figure D-27
O&M Costs for Activated Alumina with No Regeneration
1e+6
- 1e+5 -
O)
"55
O
O
1e+4 -
1e+3
1
2
3
0.001
0.01 0.1
Average Flows (mgd), x
Cost Equation
1.y = 16234+ 18256.8X
R2 = 0.99
2. y = 8729 + 296900x - 948889X2
R2 = .99
3. y = 31479 + 14869x + 0.0089X2
R2 = .99
Assumed pH
1.pH = 8.0 (treat 3000 BVs)
2. pH = 7.5 (treat 7000 BVs)
3. pH = 7.0 (treat 16500 BVs)
D-27
-------
Figure D-28
O&M Costs for Activated Alumina for 500 BV
1e+8
O
O
1e+7 ,
1e+6 -
1e+5 ^
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 16234+ 18256.8X
1. y = 8729 + 296900X - 948889X2
R2 = 99
3. y = 31479 + 14869X + 0.0089X2
R2 = .99
Applicable Flow Range
0.03-4.5 mgd
0.003 - 0.03 mgd
4.5-270 mgd
Removal Costs
c. y = 615503x^9952
d.y = 328932x-24371
4. y = 14304 + 507475X - 1684282X2
R2 = 0.99
5. y = 54515 + 96390x +10900X2
R2 = 0.99
6. y = 272544 + 263084X - 29.7X2
R2 = 0.99
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 rngd
D-28
-------
Figure D-29
O&M Costs for Activated Alumina for 2000 BV
1e+8 -T
CO
to
o
O
08
O
1e+7 •:
1e+6 -
1e+5 -
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 16234+ 18256.8X
1. y = 8729 + 296900X - 948889X2
R2 = .99
3.y = 31479 + 14869X + 0.0089X2
R2 = .99
Applicable Flow Range
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
Removal Costs
c. y = 396331 x + 16203
d.y = 104428x + 37364
4. y = 14181 + 499934X- 1650171 x2
R2 = 0 99
5. y = 42376 + 92455x - 5417X2
R2 = 0.99
6. y = 73604 + 96408X - 7.5X2
R2 = 0.99
D-29
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
-------
Figure D-30
O&M Costs for Activated Alumina for 3000 BV
1e+8
U)
•4-*
in
O
O
1e+7 -
1e+6 ,
1e+5 ,
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 16234+18256.8X
1. y = 8729 + 296900X - 948889X2
R2 = 99
3. y = 31479 + 14869X + 0.0089X2
R2 = .99
Applicable Flow Range
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
Removal Costs
c.y = 387073x +15443
d. y = 90419x +40421
4. y = 14084 + 493975X - 1623214.9X2
R2 = 0 99
5. y = 40890 + 91816x - 7495.2X2
R2 = 0.99
6. y = 58611 + 86399X - 5.02x2
R2 = 0.99
D-30
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
-------
Figure D-31
O&M Costs for Activated Alumina for 5000 BV
1e+8
CO
*-j
«
o
O
08
O
1e+7 ,
1e+6 ,
1e+5 -
1e+4 ,
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a.y = 16234 + 18256.8x
1. y = 8729 + 296900X - 948889x2
R2 = .99
3. y = 31479 + 14869x + 0.0089X2
R2 = .99
Applicable Flow Range
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
Removal Costs
c.y = 326580x+ 17528
d.y = 55978x + 48844
4. y = 13890 + 482040x -1569222x2
R2 = 0 99
5. y = 37913 + 90537x -11657x2
R2 = 0.99
6.y = 16872 + 63096x- 2.97X2
R2 = 0.99
D-31
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
-------
Figure D-32
O&M Costs for Activated Alumina for 7000 BV
1e+8
CO
to
o
O
oO
O
1e+7 ,
1e+6 ,
1e+5 ,
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 16234+ 18256.8X
1. y = 8729 + 296900X - 948889X2
R2 = .99
3. y = 31479 + 14869x + 0.0089X2
R2 = .99
Applicable Flow Range
0.03-4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
Removal Costs
c. y = 31 1910.7X + 17188
"
4. y = 13738 + 472723x - 1 527071 x2
R2 = 0.99
5. y = 36362 + 90167X - 14178.1X2
R2 = 0.99
6. y = 28350 + 58931x - 2.4x2
R2 = 0.99
D-32
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 • 270 mgd
-------
Figure D-33
O&M Costs for Activated Alumina for 10000 BV
1e+8
U)
O
O
1e+7 -
1e+6 ,
1e+5 -
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a.y = 16234 + 18256.8x
1. y = 8729 + 296900X - 948889X2
R2 = 99
3. y = 31479 + 14869x + 0.0089X2
R2 = .99
Applicable Flow Range
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
Removal Costs
c.y = 272957x + 18136
d. y = 46103x +47373
4. y = 13511 + 458748X - 1463845x2
R2 = 0.99
= .
5. y = 34037 + 89614x - 17960X2
R2 = 0.99
51990x-1.8x2
.
R2 = 0.99
D-33
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
-------
Figure D-34
O&M Costs for Activated Alumina for 16500 BV
1e+8
in
to
o
O
1e+7 -
1e+6 -.
1e+5 -.
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 16234+ 18256.8x
1. y = 8729 + 296900X - 948889X2
R2 = .99
3. y = 31479 + 14869x + 0.0089x2
R2 = .99
Applicable Flow Range
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
Removal Costs
c. y = 266125x +17524
d.y = 42678x + 47126
4. y = 13338 + 448159x - 1415947.5x2
R2 = 0.99
5. y = 32902 + 88917x -18685.8X2
R2 = 0 99
6.y = 17146 + 49346x-1.2x2
R2 = 0.99
0.03 - 0.086 mgd
0.36 - 4.,'5 mgd
0.003-0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
D-34
-------
Figure D-35
O&M Costs for Activated Alumina for 25000 BV
1e+8
CO
"So
o
O
ofi
o
1e+7 -
1e+6 -
1e+5 -
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 16234+ 18256.8X
1. y = 8729 + 296900X - 948889X2
R2 = .99
3. y = 31479 + 14869x + 0.0089X2
R2 = .99
Applicable Flow Range
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
Removal Costs
c.y = 242020x +18008
d.y = 37787x +46927
4. y = 13110 + 434124x - 1352455X2
R2 = 0.99
5. y = 31399 + 87994x - 19648X2
R2 = 0 99
6. y = 13015 + 45327x - 0.96X2
R2 = 0.99
D-35
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
-------
Figure D-36
O&M Costs for Activated Alumina for 50000 BV
1e+8
en
"55
o
O
1e+7 -
1e+6 -
1e+5 ,
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 16234+ 18256.8X
1. y = 8729 + 296900X - 948889X2
R2 = .99
3. y = 31479 + 14869x + 0.0089X2
R2 = .99
Applicable Flow Range
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
Removal Costs
c.y = 261694x + 15123
d. y = 35056x + 46593
4. y = 12240 + 380671x - 1110631X2
R2 = 0.99
5. y = 30204 + 88142x - 21008X2
R2 = 0 99
6. y = 10380 + 43106x - 0.68x2
R2 = 0.99
D-36
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
-------
Figure D-37
Capital Costs for Ion Exchange
1e+9 ,
1e+4
0.01
1 10
Design Flows (mgd), x
100
1000
Accessories Costs
a. y = 2x106x-199382
b.y = 106642x+576936
1. y = 11187 + 255659X - 322917X2
R2 = 1.0
2.y = 341596 + 319324X + 22658X2
R2 = .99
3. y =601203 + 104094X + 12.16X2
R2 = .99
Applicable Flow Range
0.1 -0.27 mgd
1-10 mgd
0.01 -0.1 mgd
0.27-1.0 mgd
10-430 mgd
Removal Costs
c. y = 5690918x-490868
d.y = 432316x+1500732
4. y = 32399 + 272990X + 2989548X2
R2 = 0.99
5. y = 521307 + 2138293x - 726552X2
R2 = 0.99
6. y = 1891169+ 393272X
R2=0.99
D-37
Applicable Flow Range
0.1 - 0.27 mgd
1 -10 mgd
0.01 - 0.1 mgd
0.27-1.0 mgd
10-430 mgd
-------
Figure D-38
O&M Costs for Ion Exchange for 300 BV
1e+7
1e+6 -
o
O
O
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 52578X + 3048
b. y = 56152x-2819
1. y = 2169 + 121102X- 1351017
R2 = .99
2. y = 4346 + 37861X - 4484X2
R2 = .99
3. y = 166473 + 18563X - 6.987X2
R2 = .99
Removal Costs
c. y =12018 + 33140X
4. y = 4117 + 457302X - 5359557X2
p2 _ gg
6. y = 52641 + 24123x - 2.274X2
R2 = .99
Applicable Flow Range
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
D-38
-------
Figure D-39
O&M Costs for Ion Exchange for 500 BV
1e+7
1e+6 -
o
o
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a.y = 52578X + 3048
b. y = 56152x-2819
1. y = 2169 + 121102X- 1351017
R2 = .99
2. y = 4346 + 37861x - 4484X2
R2 = .99
3. y = 166473 + 18563x - 6.987X2
R2 = .99
Removal Costs
c. y = 7499 + 19789.4X
4. y = 3427 + 236042x - 2683967X2
R2 = .99
6. y = 31335 + 14499x - 1.452X2
R2 = .99
Applicable Flow Range
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
D-39
-------
Figure 3-40
O&M Costs for Ion Exchange for 700 BV
1e+7
1e+6 -
o
O
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 52578X + 3048
b. y = 56152x-2819
1. y = 2169 + 121102X - 1351017
R2 = .99
2. y = 4346 + 37861X - 4484X2
R2 = .99
3. y = 166473 + 18563X - 6.987x2
R2 = .99
Removal Costs
C. y = 4897 + 14236.8X
4. y = 2613 + 135310X - 1497508X2
R2 = .99
6.y = 22382 +10356x - 1.037X2
R2 = .99
Applicable Flow Range
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
D-40
-------
Figure D-41
O&M Costs for Ion Exchange for 1500 BV
1e+7
1e+6 -
>.
>,
o
O
1e+4 -
1e+3
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a.y = 52578x + 3048
b. y = 56152x-2819
1. y = 2169 + 121102X -1351017
R2 = .99
2. y = 4346 + 37861 x - 4484X2
R2 = .99
3.y = 166473 + 18563x - 6.987X2
R2 = .99
Removal Costs
c. y = 2691 + 6553.9X
4. y = 1780 + 53121 x - 540145X2
R2 = 99
6. y = 10445 + 4833x - 0.4840X2
R2 = .99
Applicable Flow Range
0.03 - 0.086 mgd
0.36 - 4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
0.03-4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
D-41
-------
Figure D-42
O&M Costs for Ion Exchange for 2500 BV
1e+7
to-
"eo
O
O
1e+6 ,
1e+5 ,
1e+4 -
1e+3 -
1e+2
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Accessories Costs
a. y = 52578x + 3048
b. y = 56152x-2819
1. y = 2169 + 121102X- 1351017
R2 = .99
2. y = 4346 + 37861 x - 4484X2
R2 = .99
3. y = 166473 + 18563x - 6.987X2
R2 = .99
Removal Costs
c. y = 1207 + 4023x
4. y = 896.9 + 20070x - 189888X2
R2 = .99
6. y = 6267 + 2900X - 0.2904X2
R2 = .99
Applicable Flow Range
0.03 - 0.086 mgd
0.36-4.5 mgd
0.003 - 0.03 mgd
0.086 - 0.36 mgd
4.5 - 270 mgd
0.03 - 4.5 mgd
0.003 - 0.03 mgd
4.5 - 270 mgd
D-42
-------
Figure D-43
Capital Costs for Microfiltration
le+9
le+4
0.01
Equation
l.y = 1571500*x°-6882
R2 = 0.92
2.y = 527100x°-6145
R2 = 0.92
3.y = 437777*x°-4150
R2 = 0.92
4. y = 785557x - 3523699
5. y=149019x +452020
R2 = 0.99
a. y=167159x +270618
b. y = 742357x + 53818
1 10
Design Flows (mgd), x
Applicable Flow Range
1.0.01-11 mgd
2. 0.01-0.1 mgd
3. 0.27 -1 mgd
4. 11-430 mgd
5. 10-430 mgd
a. 1 - 10 mgd
b. 0.1 - 0.27 mgd
100
1000
• Removal Cost
o Accessories Cost
D-43
-------
Figure D-44
O&M Costs for Microfiltratiori
le+8
le+7 J
le+5
o
U
§ le+4
le+3 J
le+2
0.001
0.01
0.1 1 10
Average Flows (mgd), x
100
1000
Equation
l.y = 75361x°-9522
2. y = 83.5 + 67926.2x
R2 = 0.99
Applicable Flow Ranp;e
1.0.003-10 mgd
2. 10-430 mgd
3. 0.003-0.1 mgd
4. y = 32595.4x°'5378 4. 0.1 - 0.36 mgd
R2 = 0.99
5. y = 30484 + 9908.3x 5. 4.5 - 270 mgd
R2 = 0.99
a. y = 13925 + 13588.2x a. 0.36 - 4.5 mgd
• Removal Cost
o Accessories Cost
D-44
-------
Figure D-45
Capital Costs for Ultrafiltration
le+9
le+8 :
le+7 -
cj
u
le+6 -
le+5 -
le+4
0.01
1 10
Design Flows (mgd), x
100
1000
Equation
l.y=1444700*x°-6569
R2 = 0.99
2.y = 606774x°-6569
R2 = 0.99
3.y = 437777x°-4150 3.0.27-
R2 = 0.99
4. y = 694960x - 2862850
5. y = 149019x +452020
R2 = 0.99
a. y=167159x +270618
b. y = 709577x + 62669
Applicable Flow Range
1. 0.01 -18 mgd
2. 0.01 - 0.1 mgd
4. 18-430 mgd
5.10-430 mgd
a. 1 -10 mgd
b. 0.1-0.27 mgd
D-45
• Removal Cost
o Accessories Cost
-------
le+8
Cfl
3
le+7 -
le+6 -.
le+4 -.
le+3 -
0.001
Figure D-46
O&M Costs for Ultrafiltration
0.01
Equation
0.1 1 10
Average Flows (mgd), x
Applicable Flow Range
1. y = 131872.9x°'8065 1. 0.003 - 8 mgd
R2 = 0.60
2. y = -0.0023 + 86445.7x 2. 8 - 270 mgd
. = 60707.7x°-7870
R2 = 0.99
4.y = 33353.8x°'5568
R2 = 0.99
5. y = 304485 + 9908.3x
R2 = 0.99
a.y=1091-88790x
b. y= 13998 +13572. Ix
3. 0.003 - 0.03 mgd
4. 0.09-0.36 mgd
5.4.5 - 270 mgd
a. 0.03 - 0.09 mgd
b. 0.36 - 4.5 mgd
D-46
100
1000
• Removal Cost
o Accessories Cost
-------
Figure D-47
Capital Costs for Nanofiltration
le+9
le+8 -.
le+3
0.01
Equation
1 10
Design Flows (mgd), x
Applicable Flow Range
1. y = 13673 + 215457x + 17043869x2 1. 0.01 - 0.1 mgd
R2 = 0.99
2. y = 7657 + 120656x + 9544566x2
4. y
5. y
R2 = 0.99
= 437777x°-4150
R2 = 0.99
= 1271 145x- 2789165
=149019x + 452020
R2 = 0.99
6. y = 2627946x°-6285 - 667852
R2 = 0.99
a. y=167159x + 270618
b.y=818155x + 33353
c.y=1650300x + 40627
2- °-
3. 0.27 -1 mgd
4.11-430 mgd
5.10-430 mgd
6. 0.27-11 mgd
a. 1 -10 mgd
b. 0.1-0.27 mgd
c. 0.1 -0.27 mgd
D-47
100
1000
• Removal Cost
o Accessories Cost
-------
Figure D-48
O&M Costs for Nanofiltration
le+8
le+7 -:
le+6 -J
CO . . -.
- le+5
o
U
§ le+4
le+3 .=
le+2
0.001
0.01
0.1
111 i 1—i i i i 11| r—i—i i i i 111 1 1—i i i i 11
1 10 100 1000
Equation
Average Flows (mgd), x
Applicable Flow Range
1. y = 328 - 740927x 1. 0.003 - 0.3 mgd
2. y = 409950.9X0'5065 2. 0.3 - 5 mgd
3. y = 44404.5 + 183312x 3- 5 ' 27°
• Removal Cost
o Accessories Cost
4.y=101625x
R2 = 0.99
5. y = 33353.8x°-5568
R2 = 0.99
6. y = 304485 + 9908.3x
R2 = 0.99
a. y =-151.7+103241.8x
b.y= 13998+ 13572.1x
4. 0.003 - 0.03 mgd
5. 0.09 - 0.36 mgd
6. 4.5 - 270 mgd
a. 0.03 - 0.09 mgd
b. 0.36 - 4.5 mgd
D-48
-------
CO
O
U
"3
'I
O
Figure D-49
Capital Costs for Reverse Osmosis
1e+9
1e+8 -
1e+7 -
1e+6 -
1e+5
1e+4
0.01
0.1
1 • 10
Design Flows (mgd), x
Equation
a. y = 587713x+ 104772
b.y = 187950x +258005
1.y = 34570+1289740X
R2 = 0.94
2. y = 197258 + 243872X + 4825.19X2
R2 = 0.99
3. y = 7682317 + 1836106x -648.7X2
R2 = 0.99
Applicable Flow Range
1.0.01-0.1 mgd
2. 0.27 -1 mgd
3.10-430 mgd
4. 0.01-0.1 mgd
1000
• Removal Cost
o Accessories Cost
c. y = 1395912x + 240744
d. y = 2686316x-884650
4. y = 80395 + 2999395x
R2 = 0.94
5. y = 182042 + 1610996x + 8628.2X2
R2 = 0.99
6. y = 745567 + 138955x + 23.86X2
R2 = 0.99
5. 0.27 -1 mgd
6. 10-430 mgd
a. 0.1 - 0.27 mgd
b. 1 -10 mgd
c. 0.1 -0.27 mgd
d. 1 -10 mgd
D-49
-------
Figure D-50
O&M Costs for Reverse Osmosis
•
CO
O
u
1e+8 ,
1e+7 ,
1e+6 ,
1e+4 ,
1e+3
0.001
0.01
0.1 1 - 10
Average Flows (mgd), x
Equation
Removal Costs
a. y = 27350 + 641036.8X
b.y = 78049.5 + 481380.9X
°-7817
1.y = 157604x03476
R2 = 0.99
2. y = 55861 3.5x
R2 = 0.99
3. y = 580418.4 + 369743.5X
R2 = 0.99
Applicable Flow Range
1.0.003-0.03 mgd
2. 0.09 - 0.36 mgd
3. 4.5 - 270 mgd
4. 0.003 - 0.03 mgd
100
1000
• Removal Cost
o Accessories Cost
Accessories Costs
c. y = 2742 + 66502.2x
d. y = 13998 + 13572.1 x
4. y = 1965 + 92397.3x
R2 = 0.99
5. y = 33353.8X0'5568
R2 = 0.99
6. y = 30485 + 9908.3XX
R2 = 0.99
5. 0.09 - 0.36 mgd
6. 4.5 - 270 mgd
a. 0.03 - 0.09 mgd
b. 0.36 - 4.5 mgd
c. 0.03 - 0.09 mgd
d. 0.36 - 4.5 mgd
D-50
-------
APPENDIX E
ADDITIONAL CAPITAL COSTS
-------
-------
I
1
s
I!
42
-------
(9
s
1
O
s
I
at
c
Lime Soften
Table E5 -
i Flow (mgd
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Q
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0
co e
u
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4»
3
CO ^1
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CD
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5§
ft
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5
CN
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ii
si
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6 5
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Si?
CN
9 CO
[Disinfection of Finished Water
If
,_
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s?
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= s
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US
is
in" *-"
tt *-
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ss
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1 *7
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n o>
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IAddilional Fillrallon Structures
Additional Backwash Pumps
$2.772,069 I
^
f
$366.044
$203,913
lEO'ZSlt
106,635
S
K
S
CD
O
CD
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s
CN*
s
i
§
3
ft
S
9"
ft
>
r"
»
1
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IAddilional Raw Water Pumps
$5,519,520 1
00
s
$748,261
o
I
$344,021
o
0
s
i*
$122.496
o
rj
CO
0
5"
[
n
s
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1
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9
ft
J
J
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c
I
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CO
8
oc
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i
£
A
$15.401
n
n
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si"
b
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8
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1 Potassium Permanganante Feed System
$665.186 I
S
5
$337,474
$285.658
$269,338
1
S
8
4*
$203.202
$156.877
$163.271
5
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0
0
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0
4%
$82,295
9
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5
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$2,203,145 I
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1
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$925.293 i
8
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3
D
I
5
§
CN
•ft
SJ
ri
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9
•ft
ft
ft
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I
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[Land (Low Estimate)
$9,252,929 I
,_
S
^
X
o
X
M
$611.982
CN
c
I
CO
CO
en
S"
s
$158,623
$100,000
$100,000
$100,000
$100.000
$100.000
S
ft
§
»
§
i
s
5
!
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c
a
p
I
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lift
e
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(C
n
9
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Si
to
d
9
9"
g
CO,
i
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9
5
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[Permitting (Low Estimate)
u
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5
j
u
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a
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s
1
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s
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£
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$409,741
$318.893
$247,905
$207,545
§
§
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$132,903
$102,625
§
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to
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r
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u
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$1.117,179
$292,193
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Table E7 - Activated Alumina Additional Capital Costa
1
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[Backup Sodium Hydroxide Feed System
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$116.143
$77.251
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£
s
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$366,044
$203,913
$152.031
$106.635
z
49
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>,519.520 ]
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$746.261
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$122,505
$122.496
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$464.947
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1
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CN
a
it
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$877,675
$222,244
$119,188
X
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$100.000
$100.000
$100,000
$100.000
$100.000
$100.000
$100,000
3
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$100.000
[Land (High Estimate)
il
41
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218,168
124,395
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nished Water
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$422.055
273,522
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[Additional Rege
$2.772.069 i
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366,044
3
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$122,505
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$26,030
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$218,168
| $124.395
IT
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(Disinfection of Finished Water
2,772,069
•*
$1,393.594
| $366,044
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APPENDIX F
BASIS FOR REVISED ANION EXCHANGE COSTS
-------
-------
BASIS FOR REVISED ANION EXCHANGE COSTS
Very Small Systems Best Available Technology Cost Document (9/93)
• The costs in this document were used for systems with design
flows between 24,000 and 270,000 gal/day
• The equipment cost was based upon $2 10 /ft3 for resin and a
contactor cost using the following equation [source: 5/17/94
phone conversation with Jim Moore of MPI] :
contactor cost = (107 .2) (contactor volume) + 1162.6
where contactor volume = (1.5) (resin volume)
contactor cost was based on regression on tanks of
several sizes
Empty-bed Contact Time (EBCT) =2.5 min [from text p. A-135]
• The calculations for the equipment costs based on the design
flows for the first three size categories were as follows:
Flow Cat #1
Design Flow = 24,000 gal/day
Design Flow = (24,000 gal /day) (1 day/24 hr) (1 hr/60 min)
Design Flow = 16.7 gal /min
Resin Vol = (16.7 gal/min)(2.5 min)
Resin Vol = (41.75 gal) (1 ft3/7.48 gal)
Resin Vol =5.6 ft3
Resin Cost = ($210/ft3) (5.6 ft3)
Resin Cost = $1,176
Contactor Cost = (107.2) (1.5) (5.6) + 1162.6
Contactor Cost = $2,063
Total Cost = $1,176 + $2,063 = $3.2K which is MPI's listed
cost
Flow Cat #2
Design Flow = 87,000 gal/day
Design Flow = 60.4 gal/min
Resin Vol =20.2 ft3
Resin Cost = $4,242
Contactor Cost = $4,411
Total Cost = $8,652 « $8.6K which is MPI's listed cost
F-l
-------
Flow Cat #3
Design Flow
Design Flow
270,000 gal/day
187.5 gal/min
Resin Vol = 62.7 ft3
Resin Cost = $13,167
Contactor Cost = $11,245
Total Cost = $24,412 « $24.3K which is MPI's listed cost
Process schematic for Ion Exchange (Figure A-23) includes
brine storage tank, brine feed pump, salt storage, and piping
that do not appear to be included in the equipment cost --
this was confirmed in a 5/31/94 phone conversation with Jim
Moore of MPI
The conversion of equipment costs to capital costs utilized a
straight percentage approach -- all indirect: costs were based
upon a percentage of the equipment cost without any minimums.
Percentages supposedly used to calculate the indirect costs
for very small systems (from p. 2-2):
Table 1. Listed Indirect Cost Percentages
Used to Derive Capital Costs
mac
i.
2.
3.
4.
5.
6.
7.
8.
Installation
Engineering
Contractor ' s Overhead &
Profit
Legal, Fiscal, & Admin
Sitework &
Interconnecting Piping
Electrical
Standby Power
Contingencies
PERCENTAGE OR COST
30% of equipment cost
10% of total cost
12% of total cost
3% of total coist
6% of total cost
$1300 per installed HP
5% of installed equip cost
0
NOTE: Total Cost = Installed Equipment Cost + Standby Power
Electrical
Percentages actually used to calculate the indirect costs for
very small systems -- a modification of large decentralized
indirect cost factors:
F-2
-------
Table 2. Actual Indirect Cost Percentages
Used to Derive Capital Costs
•" ''^" >••'. ' • : j^ •'•'•' • • *' £TE»:'r;;i:'V "•• - : • • : • < • ? :• ''.'•''!
i.
2.
3.
4.
5.
6.
7.
8.
Installation
Engineering
Contractors Overhead &
Profit
Legal, Fiscal, & Admin
Sitework &
Interconnecting Piping
Electrical
Standby Power
Contingencies
l;>^r:vV\-":'-iqaua^AiiBrr]- ^'- •'.' ' '•
30% of equip cost
0%
12% of total cost
3% of total cost
15% of installed equip
cost
Not included
5% of installed equip
cost
0%
NOTE: Total Cost =
Standby Power
Installed Equipment Cost + Sitework +
• Effect of change (examined by transforming equipment cost 'x'
into capital costs):
Listed Procedure
CAP = [(1.3}x + (0.05) (1.3(X))] [1 + 0.10*+ 0.12 + 0.03 + .06]
CAP = 1.7882x
Actual Procedure
CAP= [(1.3)x + (0.05) (1.3(x)) + (0.15) (1.3(x) )] [1 + 0.12 +
0.03]
CAP = 1.794x
Thus, there is a very small increase in costs for ion exchange
using the actual procedure. Electrical is not a separate cost
component in the Very Small System BAT Cost Document.
Revised Anion Exchange Costs -- Basis
• The basis for the revised anion exchange costs for systems
with design flows in the range between 24,000 and 270,000 are
discussed first since the most extensive revisions occurred
F-3
-------
for these costs. The revisions discuss changes that were made
in the costs for both very small systems and large,
decentralized systems over the plant size range.
Cost Bases for Models (Year);
• WATER Model -- December 1983
• Very Small Systems BAT Cost Document -- First Quarter
1992
• Phase VI-B IOC Technology & Cost Document -- December
1991
The Phase VI-B cost indices will be used to update the WATER
model costs to the same cost year basis as the Very Small
Systems BAT Cost Document. The cost modifying factors for the
major components of the construction cost for anion exchange
are derived below using Bureau of Labor Statistics (BLS)
indices:
Description 12/83
General Purpose
Machinery (BLS 114) 309.3
Pipes & Valves
BLS 1149 327.0
12/91
390.9
447.2
Factor
1.26
1.37
In "Estimation of Small System Water Treatment Costs" (also
known as the WATER model), Table 62 contains construction cost
for anion exchange based on various total resin volumes. The
relevant portions are summarized below:*
Table 3. Relevant Cost Components from WATER Model
Cost
Category
Manufactured
Equipment
Pipes & Valves
Stafcai Resin Volume {ft*}
: ' V
$3,100
$800
1? '1
$8,600
$800
S4
$23,100
$1,000
188
$64,100
$2,600
The updated costs (12/91 basis) for these components using the
factors that were derived above are summarized below:
F-4
-------
Table 4. Updated Relevant Cost Components from WATER Model
cost
category
Manufactured
Equipment
Pipes & Valves
TOTAL
Total Resin volume {ft*}
4
$3,900
$1,100
$5,000
17 i
$10,800
$1,100
$11,900
i 54
$29,100
$1,400
$30,500
188
$80,800
$3,600
$84,400
From the calculations that were previously performed to
determine resin volume, the associated resin volume for each
design flow is as follows:
Table 5. Resin Volume versus Design Flow for
First Three Size Categories
-. ; V • '&^^^*3^^v^i
24,000
87,000
270,000
is^^aA^
5.6
20.2
62.7
Interpolating for costs:
Size Cat #1
(5.6 - 4}/(17 - 4) = (x)/($ll,900 - $5,000)
(13)(x) = 11,040
x = 850
Equipment Cost = $5,000 + $850 = $5,850 [Design Flow = 24 kgpd]
Installed Equipment Cost = (1.3)($5,850)
Installed Equipment Cost = $7,600
Size Cat #2
(20.2 - 17)/(54 - 17) = (x)/($30,500 - $11,900)
x = $1,600
Equipment Cost = $11,900 + $1,600 = $13,500
Installed Equipment Cost = $17,600
F-5
-------
Size Cat »3
(62.7 - 54)7(188 - 54) = (x)/($84,400 - $30,500)
x = $3,500
Equipment Cost = $34,000
Installed Equipment Cost = $44,200
The equipment costs for several actual existing ion exchange
plants was examined to determine if the revised estimates for
equipment cost were more accurate than the original estimates using
the cost equations. An estimate of the design flow is also
provided with the cost estimates. This table includes costs for
both cation and anion exchange units. The cation exchange cost
estimates for existing plants were converted to anion exchange
costs by using $210/ft3 for the resin cost instead of $70/ft3. The
resin costs are taken from the WATER Model. The $210/ft3 cost for
anion exchange resins and the $70/ft3 for cation exchange resins
were also used by Malcolm Pirnie to generate the initial point
estimates for anion and cation exchange that were regressed to
develop the cost equations. The type of exchange unit is
identified along with the location of the plant. The installed
anion exchange -equipment costs estimated using the equations in the
Very Small Systems Document and using the revised approach are
included for direct comparison with the actual costs. The costs
generated by the revised approach track much better with the actual
costs from existing treatment facilities.
F-6
-------
Table 6. Actual and Estimated Costs of Ion Exchange Plants
Source
Quail Creek
(CX -> AX)
Jefferson Co.
(AX)
Orig Appr (AX)
Size Cat #1
Rev Appr (AX)
Size Cat #1
Orig Appr (AX)
Size Cat #2
Rev Appr (AX)
Size Cat #2
Redhill Forest
(CX -> AX)
Orig Appr (AX)
Size Cat #3
Rev Appr (AX)
Size Cat #3
Plant Sl««s
'
14.4
14.4
(estimate)
24
24
87
87
180
270
270
installed
Equip Cost {$}
$6,620
$5,900
$4,200
$7,600
$11,200
$17,600
$27,170
$24,300
$44,200
cost Basis
{Yea*}
1991
1987
1991
1991
1991
1991
1985
1991
1991
The next step was to transform the equipment costs into
capital costs using the indirect cost factors that are listed below
the following table. The installed equipment costs for the first
three size categories have already been calculated. The basis for
the indirect costs is discussed in the April 15, 1994, memorandum
from Mary Sands, Chris Lough, and Shirley Smith of DPRA to Ben
Smith and Jeff Kempic of EPA. This memorandum reviewed and
compared the design and operating assumptions between the three
water treatment cost models for consistency. The recommended ideal
design and operating assumptions were developed for each existing
cost model for electricity, standby power, sitework/interconnecting
plumbing, legal/fiscal/administrative, engineering, contractor's
overhead and profit, and contingencies and are summarized below the
following table.
F-7
-------
P-8
-------
The cost transformation factor for calculating the costs for
the building, road, and fence around the building were obtained
from Chapter 4 in the "Very Small Systems BAT Cost Document" on
additional costs. The key design assumptions are as follows: the
process area is square, there is a 15-foot buffer between the
process area boundary and the site boundary, the fence follows the
site boundary, and a 75-foot turnaround is located at the end of
the road. These costs would only be applied to systems that do not
have any treatment installed. The cost components for a building
are: basic storage building, foundation, electrical wiring, HVAC,
and plumbing. For very small systems, a pre-fabricated building
installed on a concrete slab was assumed with a cost factor of
$40/ft2. Access roads for very small systems are assumed to be
gravel with a cost factor of $0.90/ft2. The costs for a 6-foot
chain-linked fence are $11.95/foot with $346 for four corner posts
and $800 for a 20-foot swinging gate. For ground water systems
that only disinfect, only a new building will be assumed to develop
costs -- the road and fence will be assumed to be adequate since
there is only a minor increase in costs that will be accounted for
in the cost of a new building.
The costs derived for very small systems in the first three
size categories using the cost equations for anion exchange in the
"Very Small Systems BAT Cost Document" and the two revised
approaches are presented in the following table.
Table 8. Costs Derived Using the Three Mechanisms for
the First Three Size Categories
Design: Flow
(kgpd)
24
87
270
**ry"fia*tii
Systems DOC. .
5.7 k$
10.4 k$
28 k$
Revised Appr,
No Treatment
31.6 k$
46.5 k$
86.7 k$
Revised Appr.
Disinfct* only
26.5 k$
41.3 k$
81.4 k$
Since the building, road, and fence are listed as additional
costs in the "Very Small Systems BAT Cost Document," they are not
a component of the cost equations for each technology. However,
since they are included as a component of the revised costs, a more
direct comparison of the costs generated by each model would
include these costs in the Very Small Systems Document costs. This
is presented below along with the capital costs that were generated
using the WATER model (from the 9/93 Treatment and Occurrence
Document):
F-9
-------
Table 9. Comparison of Costs Derived Using Very Small Systems
Approaches with Costs from the WATER Model
Besigst Flosr
{fcspa>
24
87
270
V8& B0«
W/ Aujld. Costs
15.7 k$
21 k$
40.2 k$
Eevisea x$pr«
No Treatment
31.6 k$
46.5 k$
86.7 k$
WATER
Model
25 k$
40 k$
71 k$
The WATER Model costs from the Arsenic Treatment and
Occurrence Document were examined and found to be incorrect. The
errors could not be isolated, but a comparison with the anion
exchange costs developed for uranium removal indicated that the
arsenic costs were seriously underestimated. In addition, the
construction cost curve in the WATER Model was examined and found
to be in error. Two of the six conceptual designs were plotted
incorrectly in Figure 51 and the cost curve appears to be based on
the incorrectly plotted points. From Table 62, the construction
cost for a total resin volume of 17 ft3 should be $29,600 including
design contingencies instead of the $39,000 that was plotted. The
construction cost for a total resin volume of 54 ft3 should be
$68,100 (w/o contingencies) and $78,300 (w/ contingencies) instead
of the $58,000 that was plotted. The data points for the six
conceptual designs were plotted both with and without design
contingencies being included in the estimation of the construction
costs. The plotted cost curve actually best represents a
combination of designs with and without contingencies. When the
total resin volume is less than 54 ft3, the construction costs from
the cost curve include design contingencies. When the total resin
volume is greater than 54 ft3, the construction costs from the cost
curve do not include contingencies. The revised costs for the
first three system size categories using the WATER Model were
calculated in the following manner:
SIZE CAT #1
Assuming that the EBCT = 2.5 min. (same assumption as the Very
Small Systems BAT Cost Document cost equation for anion exchange),
then the resin volume = 5.6 ft3. From Figure 51 in "Estimation of
Small System Water Treatment Costs" (WATER Model),, the construction
cost = $19,500.
The mechanism of updating the costs to 12/91 is described in
Section 4 of "Estimation of Small System Water Trciatment Costs" and
is discussed below. Each of the cost components for the
construction cost is summed for the six conceptual designs so that
weighted percentages can be determined for each cost component.
Table 10 shows the derivation of the weighted percentage for each
of the cost components.
F-10
-------
Table 10. Relative Percentages for Construction Cost Components
Construction cost
Coftpo&s&t
Excavation &
Sitework (Labor)
Manufactured
Equipment
Concrete
Steel
Labor, Installation
Pipes & Valves
Electrical
Housing
TOTAL
Sum of Six Designs
in Table 62
$22,700
$360,500
$26,400
$34,300
$56,500
$10,900
$18,600
$87,100
$617,000
Weighted Percentage
at Subtotal Costs
3.68%
58.43%
4.28%
5.56%
9.16%
1.77%
3.01%
14.12%
100%
After determining the percentage of the construction costs
that each of the components contributed, each component was updated
using the appropriate indices. The December 1983 (from original
document) and the December 1991 (from Phase VI-B IOC Document)
indices for each cost component are summarized below:
F-ll
-------
Table 11. Cost Indices Used for Updating Construction Costs
, co»stru«tio«- coat: , :
' Cer 19*3
Value of Index
368.68
309.3
317
359.5
368.68
«
327.0
243.7
356.30
ZMKMribar itaa,"
Value of Index
450.8
390.9
366.7
329.5
450.8
447.2
280.2
413.2
The updated costs are calculated below:
Updated Construction Cost Components;
Excav & Sitework = (0.0368)($19,500)(450.8/368.68) » $877
Manufactured Equip = (0.5843)($19,500)(390.9/309.3) = $14,400
Concrete = (0.0428)($19,500)(366.7/317) = $965
Steel = (0.0556)($19,500)(329.5/359.5) = $994
Labor = (0.0916)($19,500)(450.8/368.68) = $2,184
Pipes & Valves = (0.0177)($19,500)(447.2/327.0) = $472
Electrical = (0.0301)($19,500)(280.2/243.7) = $675
Housing = (0.1412)($19,500)(413.2/356.3) = $3,193
TOTAL = $23,760
Since Res Vol < 54 ft3 Design Contingencies @ 15% = NA
Updated Construction Cost = $23,760
NOTE: This value almost exceeds the $25,000 capital cost listed in
the Arsenic Document without indirect costs for contractor's
overhead and profit, sitework and interconnecting plumbing,
engineering and legal, fiscal, & administrative being added to
produce the capital cost. The procedure for eidding these other
indirect costs to derive the capital costs is detailed in Section
7 of the "Estimation of Small System Water Treatment Costs." The
percentages used for contractor's overhead & profit and engineering
that were used in Section 7 of the report were used to derive the
capital costs because the MPI report listed ranges for these
F-12
-------
indirect costs. The total construction cost was used to determine
the legal, fiscal & administrative costs from Figure 172 of the
report, even though the figure states that the subtotal of all
other construction costs should be used for this calculation. An
examination of the example cost derivations listed in Section 7
revealed that the subtotal of all construction costs was not used
to derive the legal, fiscal & administrative costs. The derivation
of the capital costs is listed below:
Updated Construction Cost = $23,760
Sitework @ 15% (MPI assumption) = $3,564
Total Construction Cost = $27,324
Contractor Overhead & Profit @ 12% (MPI: 8.5 - 12%) = $3.279
Subtotal = $30,603
Engineering @ 10% (MPI: 8.5 - 12%) = $3,060
Subtotal = $33,663
Legal, Fiscal & Administrative (from Figure 172) = $2,900
CAPITAL COST = $36,563
SIZE CAT #2
Using Figure 51 with a resin volume of 20.2 ft3, the
construction cost = $36,500
Updated Construction Cost Components
Excavation & Sitework = $1,639
Manufactured Equipment = $26,872
Concrete = $1,812
Steel = $1,867
Labor = $4,079
Pipes & Valves = $885
Electrical = $1,263
Housing = $5,978
$44,395
Design Contingencies @ 15% = NA
Updated Construction Cost - $44,395
Sitework & 15% = $6.659
Total Construction Cost = $51,054
Contractor O&P @ 12% = $6,126
Subtotal = $57,180
Engineering @ 10% = $5,718
Subtotal = $62,898
Legal, Fiscal & Admin. (Fig #172) = $5,100
CAPITAL COST = $67,998
Res Vol < 54 ft3
SIZE CAT #3
Using Figure 51 with resin volume = 62.7 ft3, the construction
costs = $68,000.
F-13
-------
Updated Construction Cost Components
Excavation & Sitework
Manufactured Equipment
Concrete
Steel
Labor
Pipes & Valves
Electrical
Housing
Design Contingencies @ 15%
Updated Construction Cost
Sitework @ 15%
Total Construction Cost
Contractor O&P & 12%
Subtotal
Engineering @ 10%
Subtotal
Legal, Fiscal & Admin (Fig 172)
CAPITAL COSTS
$3,053
= $50,063
$3,376
$3,478
$7,599
$1,649
$2,359
= $11.138
$82,715
= $12.407
= $95,122
= $14.263
= $109,390
= $13,127
= $122,517
= $12.252
= $134,769"
= $8.300
= $143,069
Res Vol > 54 ft3
The previous table that compared the revised costs with the
costs produced by the cost equations with the additional costs and
the initially calculated WATER Model costs has been updated. The
revised WATER Model costs are considerably higher than the
previously listed costs using the revised approach. The costs
developed using the revised approach are lower than the actual
costs using the WATER Model, but they are higher than the costs
previously calculated using the cost equations and the WATER Model
that are presented in the Arsenic Treatment and Occurrence Document
(9/93).
Table 12. Comparison of Costs Derived Using Very Small System
Approaches with Revised Costs from the WATER Model
Design* Flow
{fcspSJ
24
87
270
•'' nw'atxs
W/-A&3U £Q*tS *
15.7 k$
21 k$
40.2 k$
Sevised Appr,
m> Treatment
31.6 k$
46.5 k$
86.7 k$
< . WA3rJ5**
Model
36.6 k$
68.0 k$
143 k$
The costs that were derived using the revised approach were
compared with estimates that were received from Culligan on anion
exchange to remove nitrate. Culligan provided a cost estimate for
a modular unit to treat water for each of the first four size
categories. The design flow was not specified, so there may be
some differences between the bases for the two cost estimates.
However, the two sets of estimates compare very favorably and the
F-14
-------
Culligan data further confirm the use of the revised approach for
anion exchange capital costs. Table 13 contains the costs
generated by the revised approach and the cost estimates supplied
by Culligan.
Table 13. Comparison of Capital Costs Generated using the
Revised Approach with Estimates Supplied by Culligan
Population Size
Category
25 - 100
101 - 500
501 - 1,000
Revised Approach
Cost Estimates
31.6 k$
46.5 k$
86.7 k$
CttlliQfaa Cost
Estimates
33 k$
42 k$
66 k$
The large, decentralized system costs have also been revised.
The equipment costs are the same as were derived previously for the
very small systems. The indirect costs that were previously used
in the "Very Small Systems BAT Cost Document" were as follows:
Table 14. Indirect Cost Percentages .Used to Derive Capital Costs
ITEM
1.
2.
3.
4.
5.
6.
7.
8.
Installation
Engineering
Contractors Overhead &
Profit
Legal, Fiscal, & Admin
Sitework &
Interconnecting Piping
Electrical
Standby Power
Contingencies
PEHCESTAGE
30% of equip cost
15% of complete total
cost
12% of total cost
3% of total cost
15% of installed equip
cost
Not included
5% of installed equip
cost
20% of total cost
NOTE: Total Cost =
Standby Power
Installed Equipment Cost + Sitework +
Complete Total Cost = Total Cost + Contractor Overhead
and Profit + Legal and Administrative
P-15
-------
VLIZED SYSTEMS
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F-16
-------
The recommended ideal design assumptions for the indirect cost
factors that were used for very small system will also be used for
large decentralized systems. The only exceptions are buildings and
roads. For buildings, it is assumed that the buildings would have
to be architectured and fabricated on-site to comply with local
building codes or community aesthetic standards. The cost of these
buildings would be approximately $150/ft2. For roads, it was
assumed that a paved road would be required to meet local aesthetic
standards. The cost for a paved road is estimated to $1.70/ft2
based on 2 inches of binding asphalt and 1.5 inches of wearing
asphalt. These costs are summarized in the table on the preceding
page. The same assumption about the need for buildings, roads, and
fences will be applied to large, decentralized systems.
The costs derived for large, decentralized systems in the
first three size categories using the mechanism in the "Very Small
Systems BAT Cost Document" and the two revised approaches are
presented in the following table. Since the building, road, and
fence are listed as additional costs in the "Very Small Systems BAT
Cost Document," they are not a component of the cost equations for
each technology. However, since they are included as a component
of the revised costs, a more direct comparison of the costs
generated by each model would include these costs in the Very Small
Systems Document costs. The Very Small Systems Document costs
presented in Table 16 do include those additional costs.
Table 16. Costs Derived Using the Three Mechanisms for
First Three Size Categories
Design Flow
OcB&ai
24
87
270
Y-fery Small
Systems Do<2.
32.8 k$
47.6 k$
90.6 k$
Revised Appzr* '
So Treatment
46.9 k$
63.2 k$
107 k$
Revised-'. -ai^&kC;
DiBlnfct^Qn^:;
39.8 k$
56.1 k$
100 k$
The Very Small Systems Document develops costs for use up to
design flows of 270 kgpd (system size categories 1 thru 3). The
WATER Model is used to develop costs for system size category 4 ,
which has a design flow of 650 kgpd. The previously discussed
procedures for using the WATER Model were used to derive the
construction costs for the design flow of 650 kgpd and are shown
below. However, the capital costs are derived differently based
upon recommendations from DPRA. The costs derived using the DPRA
recommendations for indirect costs are compared below with the
standard procedure used in the Arsenic Treatment and Occurrence
Document.
Design Flow = 650,000 gal/day = 451 gal/min
Resin Volume = 151 ft3 — assuming EBCT =2.5 min
F-17
-------
From Figure 51, the construction cost = $112,500
Updated Construction Cost Components
Excavation & Sitework = $5,062
Manufactured Equipment = $83,076
Concrete = $5,570
Steel = $5,733
Labor, Installation = $12,600
Pipes & Valves = $2,723
Electrical = $3,893
Housing = $18.422
$137,079
Design Contingencies @ 15% = $20,562
Updated Construction Cost = $157,641
Using the assumptions from the WATER Model and the Arsenic
Treatment and Occurrence Document produces the following capital
cost:
Updated Construction Cost = $157,641
Sitework @ 15% = $23,646
Total Construction Cost = $181,287
Contractor O&P @ 12% = $21,754
Subtotal = $203,041
Engineering @ 10% = $20,304
Subtotal = $223,345
Legal, Fiscal & Admin (Fig #172) = $13,000
CAPITAL COST $236,345
Using the recommendations provided by DPRA on indirect cost
percentages and the WATER Model format produces the following
capital cost:
Updated Construction Cost = $157,641
Sitework @ 5% = $7,882
Standby Power @ 5% (min = $5,000) = $7,882
Total Construction Cost = $173,405
Contractor O&P & 15% = $26.011
Subtotal = $199,416
Engineering & 15% = $29,912
Subtotal = $229/328
Legal, Fiscal & Admin (Fig #172) = $12,500
CAPITAL COST = $241,828
The capital costs produced by the two procedures are very
similar. The revised approach produced slightly higher capital
costs and it is preferred because the same percentages are used to
derive the capital costs in the first three size categories. Thus,
the same design assumptions and the same indirect cost factors are
being used to estimate all small system costs even though two
different cost models are being used.
F-18
-------
For large systems (> 1 MGD), the WATERCO$T Model is used to
calculate capital costs. The WATERCO$T computer model was run for
the design flows that represent systems size categories #5-12.
In this model, not all of the indirect cost factors can be
controlled. The contractor's overhead and profit is selected by
the model and ranges from 8.5 - 12%. In examining the anion
exchange runs for the eight design flows, the percentage selected
decreases with increasing system size. The variables that could be
controlled directly in the WATERCO$T Model are engineering,
sitework & interface piping, and standby power. The percentages
used for each of these indirect cost factors are consistent with
those used for the other two cost models: engineering (15%),
sitework and interface piping (5%), and standby power (5%). The
printouts listing the capital costs for the eight size categories
are included in Appendix A.
The capital costs generated by the corrected WATER Model for
size categories 3 and 4 were compared with the costs from the
WATERCO$T Model for size categories 5 through 7 to examine if there
was an economy-of-scale based on the volume of treated water. The
capital costs from the WATERCO$T Model were taken from runs using
the revised indirect costs factors that were discussed above. The
capital costs were amortized over 20 years at a 7 percent interest
rate and divided by the average daily flow and by 365 days/year.
The results are shown in Table 17.
Table 17. Comparison of Amortized Capital Costs for System
Size Categories 3 through 7
System Size \
'.Ca.tego.ry •
3
4
5
6
7
Average
Daily Flow
{fcgpd}
86
230
700
2,100
5,000
Capital Costs
k$
143
242
940
2,215
4,838
ceats/fcgal
43
27
35
27
25
There is a slight increase in the amortized costs when the
basis shifts from the WATER Model to the WATERCO$T Model between
size categories 4 and 5. One key difference between the costs for
small systems and large systems is the size of the resin tanks.
The resin tanks in WATERCO$T were sized for up to 80 percent resin
expansion during backwash whereas they were only sized for up to 50
percent resin expansion in the small system cost models. The EBCT
could also be different than 2.5 minutes that was used for the
P-19
-------
small system cost models. The EBCT was not specified in the
WATERCO$T Model as a variable nor was it listed in Volume 2 of
"Estimating Water Treatment Costs." It could not be readily
determined from the design criteria in that section because the
resin cost was not provided. Thus, the expected economy-of-scale
for the capital costs will not be maintained.
The capital costs that have been derived using the three
models for the twelve standard size categories have been summarized
in Table 18. This table also lists the model that was used to
generate the capital costs. The costs presented for the first
three size categories are listed as being from the Very Small
Systems Document. These costs were produced using the revised
approach and do not represent the costs that would be estimated by
using the cost equations in the very Small Systems Document.
F-20
-------
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F-21
-------
Operating and Maintenance Costs
The operating and maintenance (O&M) costs for arsenic removal
by anion exchange were also examined to determine if they were
valid. The Very Small Systems Document presents costs based upon
daily regeneration and a regeneration requirement of 15 lb/ft3 of
resin. The standard equation is based on these regeneration
criteria. These criteria can be varied by using the other O&M cost
equation provided in the Very Small Systems Document. The
selection of the appropriate equation for determining the O&M costs
is based upon the needed regeneration frequency since the
regeneration requirement of 15 lb/ft3 is standard for all three
models. The regeneration frequency is based on the number of bed
volumes (BVs) until regeneration.
Time to Regeneration for MCL option of 2 ucr/L;
Assume breakthrough @ 300 BVs
tR = (2.5 min/BV)(300 BV)(1 hr/60 min) = 12.5 hr
where tR is the time until regeneration
Since the time until regeneration is based upon the design
flow instead of the average flow, the time to arsenic breakthrough
treating the average flow is affected by the ratio of the design
flow to the average flow. The time until regeneration for the
first three size categories is shown below.
Size Cat #1
tR » (12.5 hr)(24/5.6)
tR = 53.6 hr
Size Cat #2
tR = (12.5 hr)(87/24)
tR = 45.3 hr
Size Cat #3
tR = (12.5 hr)(270/86)
tR « 39.2 hr
The regeneration frequency of 300 BVs was based on a Malcolm
Pirnie pilot-scale study. This study did not evaluate all possible
combinations of competing anions or water quality conditions. The
Arsenic Treatment and Occurrence Document indicates that ion
exchange could be an effective removal technology when the influent
water contains less than 25 mg/L of sulfate and less than 500 mg/L
of total dissolved solids (TDS). The following table lists the
concentrations of competing anions and water quality conditions
that could affect the performance of the technology that were
present in the test water that showed breakthrough at 300 BVs.
F-22
-------
Table 19. Anion Exchange Raw water Quality Characteristics
(Results Expressed in mg/L)
Parameter
Arsenic
Selenium
Sulfate
Alkalinity (as CaCO,)
Nitrate (as N)
TDS
Free Cl,
Concentration, .
0.0409
<0.002
13.5
90
0.3
222
1.0
Thus, higher concentrations of competing anions could be present in
the raw water of systems that could choose to use ion exchange to
comply with an arsenic standard. Those systems with higher
concentrations of competing anions would likely need to regenerate
the resin more frequently than the times that have been previously
calculated. in addition, the time calculated above using the
average flow may overestimate the actual time until regeneration
since demand may be higher during some periods when the unit is in
operation. Therefore, a margin of safety needs to be incorporated
because small systems would not be able to afford to monitor when
breakthrough occurs for arsenic. Thus, costs will be prepared for
this option (2 ug/L) assuming daily regeneration.
The basis for the O&M costs assuming daily regeneration in the
Very Small Systems Document is provided in a .table on p. A-139 and
is reproduced below:
Table 20. Basis for O&M Cost Equation
O&M COSTS FOR ANION EXCHANGE
j&srtfi Plow :
teSixS* :
5.6
87
270
C&a&ic&.is
($/srtr
1,597
5,789
17,965
Power
C$/Y*}
142
166
235
t&intea,
Material
if/yr)
601
1,615
4,559
Total OScM
W/Q Labor
**/yr*
2,340
7,570
22,759
From this table, the two key components of the O&M costs are the
cost of the chemicals used for regeneration and the cost of the
maintenance material. These two cost components were evaluated to
ensure that the O&M cost equation based on these data were valid.
F-23
-------
The amount of chemicals (sodium chloride) used for
regeneration was calculated using the following equation:
Chemical Demand = (# regens/day) (15 lb/ft3) (Resin Vol) (365 d/yr)
For systems serving less than 1 mgd, sodium chloride was assumed to
cost $105/ton (from Table 3-5 of the Arsenic Treatment and
Occurrence Document). Thus, the chemical cost would be calculated
by multiplying the chemical demand by the unit cost for the
chemical.
Size Cat ftl
Chemical Demand = (1/d) (15 lb/ft3) (5.6 ft3) (365 d/yr)
Chemical Cost = (30,660 Ib/yr)($105/ton)(1 ton/2000 Ib)
Chemical Cost = $l,610/yr ~ $l,597/yr
Size Cat #2
Chemical Demand = (1)(15)(20.2)(365)
Chemical Cost = (110,595)(105)/(2000)
Size Cat #3
= 110,595 Ib/yr
= $5,806/yr « $5,789/yr
Chemical Demand = (1)(15)(62.7)(365) ^ 343,283 Ib/yr
Chemical Cost = (343,283)(105)/(2000) = $18,022/yr « $17,965/yr
The basis for the maintenance material costs in the Very Small
Systems Document is not provided, but it is assumed that the WATER
Model was used. In the WATER Model, maintenance material required
includes periodic make-up of resin, and miscellaneous replacement
components for the exchanger, the valving, and the brine tank.
Resin loss was estimated to 10%/yr for daxly regeneration and
complete replacement of the resin was assumed every five years.
Using the assumptions about resin loss and resin replacement, the
resin maintenance material costs were calculated for each of the
three systems size categories and the first four conceptual designs
in the WATER Model so that comparisons with the total maintenance
material costs could be performed. The resin replacement costs are
shown below:
Resin
Resin
Cost/yr
Cost/yr
Resin Cost/yr =
Resin Cost/yr =
Resin Cost/yr =
Resin Cost/yr =
Resin Cost/yr =
Resin Cost/yr =
Resin Cost/yr =
[(Resin Vol)(0.10) + (1/5)(Resin Vol)][$210/ft3]
(0.3)(Resin Vol)($210/ft3)
(0.3) (4) (210) = $252
(0.3) (5.6)(210) = $601
(0.3)(17)(210) = $1071
(0.3)(20.2)(210) = $1,273
(0.3) (54)(210) = $3,402
(0.3)(62.7)(210) = $3,950
(0.3)(188)(210) = $11,844
A comparison of the resin costs with the total maintenance material
F-24
-------
costs from the WATER Model for the conceptual designs and the Very
Small Systems Document for the three size categories is shown in
the table below. The difference between the two costs are also
shown in the table. The maintenance material costs for the three
system sizes are consistent with the maintenance material costs
used in the WATER Model. The two major components of the O&M costs
were confirmed during review, so the cost equations presented in
the Very Small Systems Document can be used.
Table 21. Comparison of Resin and Maintenance Material Costs
from WATER Model and Very Small Systems Document
Resin vol
*£t*>
4
5.6
17
20.2
54
62.7
188
Resin cost
($>
252
353
1,071
1,273
3,402
3,950
11,844
Main Mat $
(WAfER)
400
—
1,400
_ _
4,000
— —
13,300
Main Mat $
{V£S DOC}
_ _
601
—
1,615
—
4,559
--
Difference
£$}
148
248
329
342
598
609
1,456
The Very Small Systems Document has two cost equations that
can be used to calculate the O&M costs for very small systems. The
standard equation is based upon daily regeneration and a
regeneration requirement of 15 lb/ft3 and the other equation allows
these two components to be varied. The variable equation will be
used, so that the costs can be directly compared with the 10 ug/L
option, which will require less frequent regeneration. The labor
required for O&M is also a variable. The text of the Very Small
Systems Document indicates that equipment manufacturers estimate
O&M labor requirements to be 4 hrs/day for resin regeneration,
routine equipment maintenance, and process oversight. The cost
equations are based on 5 days/week for labor cost calculations.
However, the O&M costs were calculated based on 7 hours/week
instead of 20 hours/week. The 7 hours/week came from the
McFarland, CA anion exchange plant which required 1 hr/day of
operator attention since the process is automated. The anion
exchange plant in McFarland treats 1 mgd. This 1 hr/day
requirement based on 7 days/week produced the labor requirement
that was used to estimate costs from the cost equation for very
small systems.
O&M COSTS
F-25
-------
Size Cat #1
The O&M cost equation from the Very Small Systems Document is:
OM = 53.86[AVGr°'273 + 0 .173 [RD] [FRQ] [DBS] [AVG}'1
+ 2 09. 8 [LAB] [AVG]'1
WHERE: OM = O&M costs, cents/kgal
AVG = Average Daily Flow, kgpd
RD = Regenerant Dosage, lb/ft3
FRQ = Regeneration frequency, times/week
LAB = Available labor for O&M, hours/week
OM = 53. 86 [5. 6]-°'273 + 0.173 [15] [7] [24] [5 . 6]'1 -f 209. 8 [7] [5. 6] '*
OM = 374 cents/kgal
OM = ($3.74/kgal)(5.6 kgal/day)(365 day/yr) = $7,645/yr
Size Cat #2
OM * 53. 86 [24]-°'273 + 0 .173 [15] [7] [87] [24]-1 + 209 . 8 [7] [5 . 6T1
OM = 150 cents/kgal -> $13,140/yr
Size Cat #3
OM = 53. 86 [86]-°-273 + 0.173 [15] [7] [270] [86]'" + 209 . 8 [7] [86] -1
OM = 90 cents/kgal -> $28,250/yr
For systems size category 4, the WATER Model is used to
determine the O&M costs. The same calculation for time to
regeneration is performed to determine which cost curve will be
used to estimate costs.
Size Cat #4
tR = (12.5 hr)(650/230) » 35.3 hr
The same assumption about daily regeneration will be applied
to systems in this size category and was applied to the smaller
systems. The cost curves for daily regeneration were used to
estimate costs for maintenance material and process energy. The
building energy is not dependent upon regeneration frequency. The
unit cost for energy is $0.086/kwh and is taken from the Arsenic
Treatment and Occurrence Document. The labor requirement was not
estimated using Figure 53 and daily regeneration because the
McFarland data were used for all small systems. Using Figure 53
and a total resin volume of 151 ft3 results in 215 hours of labor
per year. Using the McFarland assumption about labor hours per
year (365 - 1 hr/day) actually increases the O&M costs for the
systems in this size category.
From Figure 52:
F-26
-------
Energy = 250 kwh/yr (Process) + 4500 kwh/yr (Building)
Energy = 4750 kwh/yr
Energy Cost = (4750 kwh/yr)($0.086/kwh)
Energy Cost = $408.50/yr
Maintenance Material = $ll,000/yr
The salt requirement is calculated using the same procedure that
was used to verify the cost equations. From the WATER Model, the
regenerant dosage is 15 lb/ft3 (p. 152) and it was determined above
that daily regeneration would be required.
Salt Demand = (1 regen/day) (15 lb/ft3) (151 ft3) (365 day/yr)
Salt Demand = 826,725 Ib/yr
Chemical Cost = (826,725 Ib/yr)($105/ton)(1 ton/2000 Ib)
Chemical Cost = $43,403/yr
Labor = 365 hr/yr @ $14.50/hr = $5,293/yr
TOTAL O&M = $408/yr + $ll,000/yr + $43,403 + $5,293/yr
TOTAL O&M = $60,104/yr -> 72 cents/kgal
LARGE SYSTEMS (> Imgd)
The O&M costs for large systems are estimated using the
WATERCO$T Model. The basis for the WATERCO$T Model for systems >
1 mgd is Volume 2 of the 1979 "Estimating Water Treatment Costs"
Document. The runs of the WATERCO$T Model that were used to
estimate the capital costs also list O&M costs and are found in
Appendix A. The O&M costs from the WATERCO$T runs were lower than
the costs that were estimated in the Treatment and Occurrence
report. This difference is due to the fact that the O&M costs in
the report that were generated for 300 BVs were the result of
multiplying the 1600 BV estimate by 5 to reflect the more frequent
regeneration. This estimate is inaccurate because several
components are not dependent upon resin regeneration -- complete
replacement of resin every five years would not change and labor
would not increase by a factor of five since the regeneration
process can be automated.
It was assumed that the O&M costs generated by the WATERCO$T
Model were based upon daily regeneration. The WATERCO$T Model does
not list the regeneration frequency as a variable. Volume 2 of
"Estimating Water Treatment Costs" also does not provide enough
information to determine the regeneration frequency that serves as
the basis for the costs. An examination of the assumptions for the
O&M costs revealed that the labor assumption for anion exchange is
inaccurate. The labor requirements were estimated based on
filtration plants and filter pumping facilities of comparable size.
This leads to a significant overestimate in the costs because anion
exchange processes can be automated and would not require the same
degree of oversight as a filtration process. An examination of the
impact of this assumption indicates that WATERCO$T would estimate
F-27
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2097 hours of labor would be required for the anion exchange
process to treat 700 kgal/day. This translates to $58,716 per year
for labor, assuming a labor rate of $28/hr. This labor rate is
based upon union scale laborers. The hourly rate is based on an
average salary of $18.74/hr including fringe benefits and a 50
percent labor overhead cost.
A revised estimate of the labor needed for operation and
maintenance was made based on the 1 mgd plant at McFarland, CA.
This plant required 1 hr/day of O&M and this was used for small
systems. For large systems, it was assumed that 10 hr/week of
labor would be required for O&M since the process would operate
full-time and could require somewhat more operator attention. The
amount of labor required per year is 520 hours, which translates
into $14,560/yr.
Thus, the calculated O&M costs were revised for size category
5 by subtracting $58,716 and adding $14,560. The O&M costs for
size category 5 are $179,984/yr. The cost per thousand gallons was
calculated by dividing by 365 and by 700 kgal/day and was 70
cents/kgal. There is little difference between the O&M costs
(cents/kgal) calculated for size category 4 (WATER Model) and size
category 5 (WATERCO$T Model). One contributing factor is the labor
rate: $14.5/hr for systems <= 1 mgd and $28/hr for systems > 1
mgd. The assumption of additional labor requirements for large
systems also adds 156 hours of labor per year for large systems.
This additional labor likely would only impact the size category 5
and 6 costs to any degree. The O&M costs for the other seven size
categories were estimated in the same manner. The O&M costs for
all twelve size categories are shown in Table 22.
The total production costs were calculated using the amortized
capital costs and the O&M costs and are expressed in cents per
thousand gallons (C/kgal). The capital costs were amortized using
a 7 percent interest rate and a twenty year loan. Table 23
contains the amortized capital costs, O&M costs and the total
production costs for the twelve size categories. The economy-of-
scale is not maintained between size categories 4 and 5. The bulk
of the increased cost between size category 4 and 5 is due to the
capital costs. The factors that caused both the capital and O&M
costs to be higher when the WATERCO$T Model is used have been
discussed previously. One other major distinction between the
WATER Model and the WATERCO$T Model is that the WATER Model is for
package plants and the WATERCO$T Model is for fully-engineered
plants. This distinction along with the other factors is why the
economy-of-scale is not maintained in the anion exchange total
production costs.
F-28
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APPENDIX G
BASIS FOR REVISED ACTIVATED ALUMINA COSTS
-------
-------
BASIS FOR REVISED ACTIVATED ALUMINA COSTS
The following approach was taken to develop costs for activated alumina operated on a
"throw-away" basis without pH adjustment for optimal run length for small systems. The costs in
the Very Small Systems Cost Document ("Very Small Systems Best Available Technology Cost
Document") and the Water Model ("Estimation of Small System Water Treatment Costs") are
based on regeneration of the media. The costs for pH adjustment for the optimal pH range (5.5 -
6.0) were not included in the activated alumina costs in the models. Since O&M costs are based
on regeneration of the activated alumina media, the most significant changes will occur in the
estimation of these costs.
The following sources were used to develop the design assumptions and cost estimates:
The Very Small Systems Cost Document contains capital cost estimates for systems with
design flows ranging from 24,000 gallons per day to 270,000 gallons per day. The Very
Small Systems Cost Document also contains operation and maintenance (O&M) costs for
average flow ranges in this document range from 5,600 to 86,000 gallons per day.
The Water Model presents capital and O&M costs that are based on seven conceptual
designs ranging from 45,000 gal/day to 1,082,000 gal/day. These conceptual designs are
used for both capital and O&M costs.
A draft ORD report entitled "Evaluation of Arsenic Removal by Activated Alumina
Filtration at a Small Community Public Water Supply" and the December 1987 masters
report entitled "Removal of Arsenic from Bedrock Wells by Activated Alumina" were
examined to obtain full-scale data on a water system in Bow, New Hampshire. The
performance and cost for the system installed at the White Rock Water Company were
discussed in these two reports.
An August 1984 report from ORD entitled "Design Manual: Removal of Fluoride from
Drinking Water Supplies by Activated Alumina" was examined to identify key components
of the process. The Design Manual for Fluoride Removal is also based on regeneration of
the activated alumina media, so it can't be used for the entire process.
DESIGN ASSUMPTIONS AND BASIS;
CAPITAL COSTS
1. The pH will not be adjusted to operate the process at the optimal pH between 5.5 and 6.0.
The activated alumina process will be operated at the natural pH of the system to simplify
the process and avoid potential problems with lowering the pH. Two pH values will be
evaluated pH 7.0 and pH 8.0.
G-l
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[Basis: Comments from AWWA Arsenic Technical Workgroup Meeting on Treatment
Technology. EPA expressed a concern about small systems adjusting pH downward and
the AWWA Technical Workgroup agreed that there were significant risks if small systems
were adjusting pH downward (the water could become very acidic if the treatment
chemicals are overdosed). The AWWA Technical Workgroup believed that lowering the
pH would require much more oversight than most small systems would have available. In
addition, the elimination of a chemical feed system will simplify the total treatment system
oversight requirements.]
2. The pH will not need to be adjusted after the activated alumina process. The activated
alumina process should not adversely affect the finished water pH. Those systems with
high natural water pH will not need a new pH adjustment process. Those systems with
lower natural water pH values will continue to rely on the existing corrosion control
process.
[Basis: Post-treatment pH adjustment is not included in the White Rock Water Company
case study. The raw water pH of the two blended sources is approximately 7.6. The
treated water is blended with untreated water in this case study. Post-treatment pH
adjustment is also not included in the two full-scale case studies from New Hampshire in
the ORD project. The raw water pH for these case studies is approximately pH 8. Since
it does not appear that activated alumina is adversely affecting the finished water pH,
additional post-treatment corrosion control is not necessary.]
3. Empty Bed Contact Time (EBCT) = 15 minutes
[Basis: The Very Small Systems Cost Document makes this assumption. The conceptual
designs in the Water Model are based on an EBCT of 7.5 minutes (this is not stated in the
document but was derived from the conceptual designs and,verified in the Very Small
Systems Document). The April 1999 Arsenic T&C Document is incorrect when it states
that the Water Model assumed an EBCT of 15 minutes. The conceptual designs in the
1984 Design Manual for Fluoride are based on an EBCT of 7.5 minutes. The EBCT in the
White Rock Water Company case study is 18 minutes. One full-scale site is based on a
15-minute EBCT and the other is based on a 4.3 minute-EBCT.]
4. The estimated cost of the activated alumina media is $75/ft3.
[Basis: The Water Model used $44/ft3 and those costs are December 1983. Those costs
are also based on Alcoa F-l activated alumina, which is no longer produced. The cost
estimate for Alcoa F-l alumina in the draft ORD report was $133 per pound. Based on
masters report, the packed density of the activated alumina (Alcoa F-l) was 55 lb/ ft3.
Using this density, the activated alumina media cost is $73.2/ft3. The Design Manual for
Fluoride had a range of costs for Alcoa F-l activated alumina. The highest was $0.70/lb.
The density of the treatment media was listed as 50 lb/ft3, so the cost would be $35/ft3.
The POU/POE Cost Document includes an estimate of $65/ft3 and these costs are 1997.
G-2
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One of the fall-scale plants uses Alcoa CPN 28x48 (fine) alumina and it cost $1.98/lb or
$93.06/ft3 based on a packed density of 47 lb/ft3 in June 1999. The other full-scale plant
uses Alcoa DD-2 14x28 (coarse) alumina and it cost $1.01/lb in December 1998. The
cost estimate was based on an average of the data and rounding upwards slightly to
account for the longer run lengths that would be generated using the fine alumina.]
5. The bed depth assumption ranged 3 feet to 5 feet depending upon the design flow. For
the design flows lower than 0.1 mgd, the bed depth was assumed to be either 3 feet or 4
feet. For the design flows between 0.1 and 1 mgd, the bed depth was assumed to be 5
feet. The treatment media was divided into several beds to maintain a realistic bed
diameter for the largest two design flows.
[Basis: The Design Manual recommends that the bed diameter be equal to or greater than
the bed depth to prevent "wall effects". The Design Manual also states that good practice
dictates that the bed depth be a minimum of three feet and a maximum of six feet. The
Water Model assumed a bed depth of 5 feet for all conceptual designs. The designs in the
Water Model did not have bed diameters that exceeded the bed depth for many of the
conceptual designs. The Design Manual recommendation about bed diameter and bed
depth was applied to all the conceptual designs except for the smallest two design flows
(that represent non-transient non-community water systems). The bed diameter is slightly
lower than the bed depth for the first two designs, but it is not a critical flaw. The data
from the three case studies where activated alumina is being operated without
regeneration, the bed depth is greater than the bed diameter.]
6. The contactor cost has been sized based on 50% bed expansion during backwash even
though backwashing is probably not necessary. The contactor cost equation used for the
anion exchange costs in the Very Small Systems Document was used to estimate the costs.
This equation was based on a regression of tanks of several sizes:
contactor cost = (107.2)(contactor volume) + 1162.6
The contactor volume was calculated as (1.5)(media volume).
[Basis: The Water Model and the Design Manual assume 50% bed expansion to size the
contactor. Backwashing is not performed on a routine basis at any of the three small
systems operating activated alumina without regeneration. One system was originally
scheduled to backwash every four months, but kter determined it to be unnecessary.]
7. The capital costs include a redundant column to allow the system to operate while the
media is being replaced in the primary column.
[Basis: Two sets of capital costs have been developed to deal with the issue of
redundancy. A second contactor has been included in one set of capital costs, but it does
not contain any media. The columns would be in parallel and operation would switch to
G-3
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the column containing the virgin media during the media replacement in the other column.
The media was not included in the second contactor to make it easier to estimate the
O&M costs.]
8. The remaining components of the process equipment include concrete, pipe and valves
and electrical and instrumentation. A cost factor of 12.26% of the manufactured
equipment and activated alumina was used for these other construction cost components.
[Basis: The conceptual designs in the Water Model provided cost estimates for a variety
of process and construction costs. The activated alumina costs are based on an EBCT of
7.5 minutes. Since this design is using an EBCT of 15 minutes, the activated alumina
costs were doubled to estimate the contribution of each component to the construction
cost. The weighted percentage of each component was calculated for both EBCT values.
The weighted percentages for concrete, pipe and valves and electrical and instrumentation
were then expressed as a percentage of the manufactured equipment and activated alumina
costs. This was done for the data with an EBCT of 15 minutes. Concrete was 2.83%,
pipe and valves were 13.55%, and electrical and instrumentation were 9.92% of the
equipment and media costs. However, the designs in the Water Model include pipes,
valves, and instrumentation for the sulfuric acid and sodium hydroxide feed systems in the
regeneration process. Those are not needed in this design. Therefore, the pipe and valves
percentage has been reduced by a factor of 3. The electrical and instrumentation were
reduced by a factor of 5 because pH sensors and alarms are not needed and the electrical
needs are drastically reduced. The main electrical components of the activated alumina
process in the Water Model are day tank mixers and pumps used for regeneration and the
electrical immersion heater for the sodium hydroxide tank. None of these components are
needed in this design. The adjusted cost factor is 9.29%. The other process costs were
calculated by multiplying the adjusted factor by the sum of the manufactured equipment
and activated alumina costs. The sum of the other process costs and the manufactured
equipment and activated alumina costs is the total process costs.]
9. The capital costs have been estimated from the total process costs using a factor of 2.5.
[Basis: The capital costs have been estimated from the process equipment costs using the
ratio in the Guide for Implementing Phase I Water Treatment Cost Upgrades. The
process costs were assumed to be 40% of the capital costs. The breakdown of capital
costs for small systems is as follows: 40% process costs, 40% construction costs, and
20% engineering costs.]
10. The capital costs estimated in the previous step do not include the cost of a building to
house the process equipment and other add-on costs related to the site. Housing costs
were assumed for all sites. Fence and road were only assumed for those systems that have
no treatment.
G-4
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[Basis: Data from the 1995 Community Water Systems Survey were used to identify the
percentage of systems without treatment. This data analysis is described in "Geometries
and Characteristics of Public Water Systems." The data from ground water systems is as
follows:
Percentage of Ground Water Systems with No Treatment by .Size Category
Thus, the cost of a road and fence was calculated and 43% of that total was included in
the unit cost for activated alumina for systems serving between 25 and 100 people.]
11. The equation for process area, which is the key variable in estimating the costs for
buildings, roads, and fences has been based on the design flow of the system. The
equation used for process area is:
Process Area = 614.6 * Design Flow + 68
Process Area = 614.6 * Design Flow*2 + 68 (for designs with redundant column)
[Basis: The basis for the cost equations for buildings, fence and road are described in
Section 4.0 of the Very Small Systems Document. The Very Small Systems Document
provided costs as a percentage of the capital costs. The process area was calculated from
the building cost estimates by dividing by 40 (building costs were based on $40/sq ft).
The process areas for the design flows in the Very Small Systems Document were
regressed and found to best fit a linear regression. The process area for the conceptual
designs in the Water Model were also regressed and fitted to a linear regression. Process
area estimates were derived using the regression equations and the design flows for the
four smallest system size categories in Table 9 of "Small System Flows for Affordable
Technology Determinations." These flows represent the composite of all systems
regardless of source or ownership type. In addition, two smaller design flows were added
to evaluate the viability of activated alumina without regeneration at small non-transient,
non-community water systems, such as schools. The process area estimates using the
regression of the data from the Very Small Systems Document were selected to make cost
estimates for buildings, road and fence. These process area estimates were larger than the
process area estimates from the Water Model for the smaller flows (generally less than 0.1
mgd). Since this option is only likely to be viable in the smaller size categories, the higher
process area estimates were selected to be conservative. Another factor that also adds to
the conservative nature of the process area estimates is that the data are based on a design
that includes regeneration. The space devoted to chemical feeders and storage tanks for
regeneration of the media is unnecessary in this design. The process area estimates were
G-5
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compared with the minimum area needed for the contactors. The process area estimates
are greater than twice the minimum area estimates for the contactors. The process area
calculations are shown in detail in the spreadsheet file entitled "procsara.xls."]
12. The building cost to house the activated alumina system is $44.80/sq ft.
[Basis: The data from both the Very Small Systems Document arid the Water Model were
examined to develop housing costs for the activated alumina system. The housing rate
listed in the Very Small Systems Document was $40/sq ft. This rate includes the
following: basic storage building, foundation, electrical wiring, HVAC, and plumbing.
The rate is for a pre-fabricated building installed on a concrete slab. The rate is based on
data from pre-fabricated building manufacturers and the 1993 Means Building
Construction Cost Data for storage buildings multiplied by a conversion factor for small-
sized buildings. The housing costs in the Water Model vary based on size with rates
ranging from $87/sq ft for the smallest conceptual design and $50/sq ft for a facility
treating approximately 1 mgd. The only documentation on the costs in the Water Model
states that in lieu of segregating building costs into several components, the housing
category represents all material and labor costs associated with the building, including
heating, ventilating, lighting, and normal convenience outlets. However, the fabrication
and building type are unknown and, most critically, the sources used to derive the building
costs are unknown. The costs in the Water Model were developed in 1983. The housing
rate from the Very Small Systems Document was selected based on better documentation
and the use of more recent data. The ENR Building Cost Index was used to update the
costs to September 1998. For more detail, see the spreadsheet file entitled
"procsara.xls."]
13. The capital costs used to evaluate activated alumina without regeneration are based on the
design with the second column and include infrastructure costs.
OPERATION AND MAINTENANCE (O&M) COSTS
There are three major components to the operation and maintenance (O&M) costs -
replacement media, energy and labor.
14. The critical variable in calculating the replacement media costs is the run length - the
number of bed volumes until the effluent concentration exceeds a specific concentration.
For this analysis, 10% arsenic breakthrough was used. The effluent would be 10% of the
influent concentration, which is equivalent to 90% removal The; run length projections
for three natural water pH values are as follows: pH 7.0 -16,500 bed volumes until
breakthrough; pH 7.5 - 7,000 bed volumes until breakthrough; and, pH 8.0 - 3,000 bed
volumes until breakthrough.
[Basis: A model was used to make run length projections in the National Compliance
Assessment and Costs for the Regulation of Arsenic in Drinking Water, 1997 assuming
G-6
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four activated alumina batch reactors in series (see Appendix C "Analysis by M. Edwards
on Performance Basis for Activated Alumina"). The model assumed that the activated
alumina-arsenic Freundlich isotherm of Rosenblum and Clifford applies and that
equilibrium is attained in each reactor. It was noted that the isotherm was developed with
sulfate concentrations of 720 mg/L and chloride concentrations of 532 mg/L, so it
probably is conservative in predicting actual equilibrium arsenic capacity at lower sulfate
and chloride. It was also noted that the model was developed on higher influent arsenic
concentrations (> 50 /ug/L), so there may also be errors in extrapolating to lower influent
arsenic concentrations. Figure A7 in the report shows the output of the model - effluent
arsenic plotted as a function of bed volumes treated. The model predictions in Figure A7
do not address sorption kinetics and temperature considerations. A "correction factor" of
2.5 was developed to translate model results into practice. Run lengths were predicted to
be 2.5 times shorter than the model predictions. One component of the "correction
factor" was that the isotherms were generated at 25 °C whereas water treatment will
occur in practice at 15 °C. It was noted that the error of the "correction factor" could not
be assessed with any meaningful degree of accuracy. The run lengths from two older
pilot-scale studies were used to support the "correction factor." Figure A7 was used to
estimate run lengths to 10% arsenic breakthrough for influent concentrations ranging from
5 to 50 //g/L. Figure A7 was also used to estimate run lengths to 50% arsenic
breakthrough. These run lengths are presented in the spreadsheet entitled "aarunlth.xls"
along with the adjusted run lengths (derived run-lengths divided by correction factor of
2.5). These run length projections are based on water at pH 6.0.
A recently published study (1998) entitled "Pilot-Plant Trials on the Removal of Arsenic
From Potable Water Using Activated Alumina" examined arsenic removal over a pH range
from pH 6 to pH 7.5. This study was investigating compliance options for a water system
in the United Kingdom that has influent arsenic at 22 /zg/L. The desired effluent
concentration in this study was 10 /zg/L because the UK standard for arsenic is likely to be
revised from 50 to 10 jUg/L. One set of trials examined the effect of pH on arsenic
adsorption. Trials were conducted using pH values of 7.5 (source), 7.0, 6.5, and 6.0,
using hydrochloric acid to adjust the pH. The EBCT and media type were held constant
at 3 minutes and 14 x 28 mesh, respectively. The 3 minute EBCT was chosen because pH
adjustment with acid increased bed life and results were required within a reasonable time
scale. The temperature of the water was 15 °C. The run lengths presented in the paper
best approximate 50% breakthrough. These run lengths are presented in the spreadsheet
entitled "aarunlth.xls." An exponential equation was derived from the run length data. A
comparison of the actual data with the estimates for the four pH values using the
exponential equation indicated a very good fit, with one exception. The equation
underestimates the run length at pH 6.5 by almost 13%. Since the run lengths for pH 6
and pH 7 are slightly overestimated, the exponential equation was used to make run length
predictions for the pH range in the study. It was also extrapolated to make run length
predictions for raw water pH values between pH 8 and pH 9. These run length
projections were for 50% breakthrough.
G-7
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The estimated run length for 50% breakthrough with an influent concentration of 25
at pH 6 is 108,000 bed volumes using the isotherm data without applying the correction
factor whereas the run length was 43,200 bed volumes using the correction factor. The
run length for approximately 50% breakthrough with an influent concentration of 22 //g/L
is 110,500 bed volumes for the pilot-plant trial at pH 6. It was determined that the
isotherm data without the correction factor best fit the pilot-plant data. An adjustment
factor was developed using the uncorrected data. The adjustment factor is the ratio of the
run length to 10% breakthrough over the run length to 50% breakthrough at pH 6.0. Run
lengths were derived for each pH value using the exponential equation and then were
multiplied by the adjustment factor to produce run lengths to 10% breakthrough. The run
lengths to 10% breakthrough for each pH value are included in the spreadsheet entitled
"aarunlth.xls."
Data from other pilot-plant trials were compared with the run length projections to verify
the estimates. One pilot-plant study investigated the effect of media size on arsenic
adsorption by activated alumina in the Simms and Azizian paper. Mesh sizes of 14 x 28
and 28 x 48 were compared at pH 6.5 and pH 7.5. The EBCT was 6 minutes. The scale
on the data from these pilot-plant trials made it easier to determine when the outlet arsenic
concentration exceeded 2 //g/L. Since the influent concentration was 22 //g/L, a
concentration that exceeded 2 ^g/L was used to determine the number of bed volumes
treated until 10% breakthrough. These data are .summarized in the following table:
Comparison of Run Length Projections with Other Pilot-Scale Data
(Number of Bed Volumes)
6.5
39,000
15,600
17,000 *
43,000
7.5
7,000
2,800
4,000
7,500
* 2 //g/L was first exceeded at 5000 bed volumes, but the outlet arsenic concentration
dropped back below it. After 17,000 bed volumes, the concentration was always higher
than 2 ,ug/L.
The data from the other pilot-plant trials supports the decision to use the run lengths
without the use of the correction factor. The run lengths for both the 14 x 28 mesh and
28 x 48 mesh activated alumina exceed the projections using this correction factor. The
run lengths for the 14 x 28 mesh are more consistent with the projections using the
correction factor. The run lengths for the finer 28 x 48 mesh exceeded both sets of run
length projections. The run length for the finer media are more consistent with the
uncorrected projections. Since the activated alumina costs are based more on the finer
G-8
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mesh media, it is more appropriate to use the higher run lengths.
The data from one other experiment in the Simms and Azizian paper also support the use
of the uncorrected projections. One conclusion from the paper was that there a linear
relationship between EBCT and arsenic adsorption. This conclusion is based on data from
another pilot-plant trial that examined the number of bed volumes treated for EBCTs of 3,
6, and 12 minutes at pH 7.5. The number of bed volumes treated until breakthrough at 10
Mg/L (50% breakthrough) increased from 9,000 to 14,000 when the EBCT went from 3 to
12 minutes. Since a 15-min EBCT was used in this design, the uncorrected run lengths
are more appropriate than the corrected run lengths.
One final factor that would increase the run length is intermittent operation at small
systems. When a fixed bed process is operated intermittently, the sorbed ions can migrate
deeper into the pore structure of the media, thereby exposing more external surface area
to the ions in solution. This factor also supports the use of the uncorrected run lengths in
developing the O&M costs for this process.]
15. The second major component in the O&M costs is incremental labor cost. Incremental
labor is the labor associated with the additional maintenance that come with a new
process. Adjustment in staffing or shifting of activities are not included in the incremental
labor. The incremental labor for the activated alumina process without regeneration is one
hour per week. In addition, time is also included for media replacement (at 10%
breakthrough) depending upon the volume of media being replaced. Sixteen hours was
assumed for resin volumes greater than 100 cubic feet and ten hours for smaller resin
volumes.
[Basis: The labor requirements under this design are minimal. The chemical feeders have
been eliminated since the media is replaced rather than regenerated and the process is
operated at the raw water pH. If backwashing every four months is necessary, then it
could be done in the allotted time. Estimated total annual labor costs were provided for
the Bow, NH systems in "Evaluation of Arsenic Removal by Activated Alumina Filtration
at a Small Community Public Water Supply." These estimates were based on four hours
per month for routine inspections and 16 hours/cycle for filter media replacement. The
hours for activated alumina media replacement were reduced for the smaller designs
because there is just one contactor and the replacement process is simpler.]
16. The labor rate for small systems was $28/hour.
[Basis: This is the labor rate recommended as a loaded labor rate for small systems at the
EPA Technology Design Workshop and is included in the "Guide for Implementing Phase
I Water Treatment Cost Upgrades."]
17. The third component in the O&M costs is the energy costs for both the building and the
process. Housing electrical energy use is based upon an annual usage of 19.5 kwh/sq
G-9
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ft/yr, which includes lighting, ventilation, and heating. Process electrical energy was
assumed to be 440 kwh/yr. Costs were based on a rate of $0.()8/kwh.
[Basis: The Water Model was used for the electrical usage estimates. The housing
electrical use is taken directly from the O&M cost discussion on activated alumina.
Lighting is required only when the operator is present, which was assumed to be average 3
hours per day. Lighting would likely be needed less frequently under the simplified design.
The process electrical energy requirements were also taken from the Water Model.
However, they were modified significantly since regeneration is not being performed in
this design. Process electrical energy is required for day tank mixers and diaphragm
pumps to feed from the day tanks. Mixers were assumed to operate one-half hour per
day, with the diaphragm feed pumps operating 24 hours per day. The other major energy
use component was the electrical immersion heater used to maintain the sodium hydroxide
tanks above freezing (14.4 °C) and it was assumed to operate 4 hours per day and 180
days per year. Since all of these major energy use components are not included in the
design, the process energy requirements in Table 67 of the Water Model were examined
for the smallest design flow (0.045 mgd). The process energy requirements were assumed
to be one-tenth of the listed energy requirements and assumed to be constant over the
entire flow range.]
WASTE DISPOSAL COSTS
18. The primary waste is the spent activated alumina. Rates for disposal at sanitary landfills
are based on tons of material A density of 55 lb/ft3 was used to convert the volume of
activated alumina into a mass for disposal.
[Basis: The frequency of activated alumina replacement that was determined in the
calculation of the O&M costs was used to determine the yearly waste volume. The
density of activated alumina was discussed in the fourth step of this calculation. There
were three sources of data on the density of activated alumina. The Design Manual listed
a density of 50 lb/ft3 and stated that it varies with packing characteristics of media in
vessel The Masters Report listed a packed density of Alcoa F-l alumina of 55 lb/ft3. The
packed density of Alcoa CPN 28 x 48 alumina was 47 lb/ft3 in one of the case studies.
The highest density was selected for the waste disposal calculations.]
19. The disposal cost for the spent activated alumina media is $60/ton.
[Basis: The disposal rate of $60/ton is taken from "Evaluation of Arsenic Removal by
Activated Alumina Filtration at a Small Community Water Supply." The cost equation for
disposal in a nonhazardous landfill from "Small Water System Byproducts Treatment and
Disposal Document" was also examined to estimate disposal costs. The cost equation in
that document had two variables: tons of sludge requiring disposal and transportation
distance. The transportation distance could range from 5 to 50 miles. Using the upper
G-10
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bound of 50 miles, the equation simplified to a multiplier of $45.58/ton. The higher
$60/ton was selected for the cost analysis because it was for disposal of spent alumina and
was for a similar volume of waste.]
20. The only other potential waste stream is backwash water. Due to the infrequent need for
backwashing and the low volume of waste, no costs were assigned for backwashing.
[Basis: Both the Very Small Systems Document and the Water Model provided a
backwash rate of 8 - 9 gpm/sq ft and a backwash duration of 10 minutes. These data were
used along with the activated alumina volume and depth calculations from the capital costs
to estimate the backwash volume. The annual volume for disposal was calculated by
multiplying volume per backwash by three (assumed every four months). The annual
volumes for even the largest systems never exceeded 20,000 gallons per year. For the
systems with design flows below 0.1 mgd, the annual backwash volume was below 3,000
gallons. An extremely small septic system could be installed for the backwash water if
necessary. Backwashing was not performed in the Bow, NH case study probably due to
the frequent media replacement. Backwashing is not performed on a routine basis at the
other two full-scale systems. One of the full-scale plants was originally scheduled to
backwash every four months, but this was discontinued after it was found to be
unnecessary. Since backwashing is not essential to operation of these plants and if
performed, produces a very small waste stream, -waste disposal costs are not significant.
Backwash water disposal costs, if needed, are covered by the conservative assumptions
that have been used throughout the costing process.]
G-ll
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APPENDIX H
REGIONALIZATION COST
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