United States

Environmental Protection

Agency

Office of Water

(4606)

EPA 815-R-00-028

December 2000

www.epa.gov/safewater

Technologies and

Costs for Removal of

Arsenic from Drinking

Water
.
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.

December 2000

INTERNATIONAL CONSULTANTS, INC.

4134 Linden Avenue

Dayton, Ohio 45432

MALCOLM PIRNIE, INC.

432 North 44th Street, Suite 400

Phoenix, Arizona 85008

THE CADMUS GROUP, INC.

135 Beaver Street

Waltham, Massachusetts 02452

 

Under Contract with the USEPA No. 68-C6-0039

Delivery Order 13
.
ACKNOWLEDGMENTS

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.

Technical 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 Cadmus Group, Inc. provided assistance in updating the

document to reflect the final rule. The Project Manager was Ron Braun of International Consultants.

The International Consultants technical support team was led 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 led by Zaid Chowdhury with project support from Anne Jack, Amparo

Flores and Spyros Papadimas. Ian Kline led the Cadmus Group team. Additional technical assistance

and consultation were provided by: Gary Amy, University of Colorado - Boulder; Joe Chwirka,

CH2MHill - Albuquerque; Dennis Clifford, University of Houston; Marc Edwards, Virginia

Polytechnic Institute; Janet Hering, California Institute of Technology; Jeff Robinson, IndianaAmerican Water Company; and Mark Weisner, Rice University.
.
i

TABLE OF CONTENTS

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 ........................... 2-8

2.2.4 Enhanced Coagulation ...................................... 2-9

2.2.5 Lime Softening ........................................... 2-10

2.3 Adsorptive Processes ............................................ 2-14

2.3.1 Activated Alumina ........................................ 2-15

2.4 Ion Exchange ................................................... 2-20

2.4.1 Introduction .............................................. 2-20

2.4.2 Effect of pH ............................................. 2-20

2.4.3 Effect of Competing Ions ................................... 2-20

2.4.4 Resin Type .............................................. 2-22

2.4.5 Process Configuration ...................................... 2-22

2.4.6 Secondary Effects ......................................... 2-24

2.4.7 Resin Fouling ............................................ 2-24

2.4.8 Regeneration ............................................. 2-24

2.4.9 Regenerant Reuse and Treatment ............................. 2-25

2.4.10 EBCT .................................................. 2-26

2.4.11 Field Studies ............................................. 2-26

2.4.12 Typical Design Parameters .................................. 2-27

2.5 Membrane Processes ............................................ 2-27

2.5.1 Introduction .............................................. 2-27

2.5.2 Important Factors for Membrane Performance ................... 2-29

2.5.3 Arsenic Removal with Membrane Processes .................... 2-30

2.5.4 Microfiltration ........................................... 2-30

2.5.5 Ultrafiltration ............................................ 2-31

2.5.6 Nanofiltration ............................................ 2-33
ii

2.5.7 Reverse Osmosis ......................................... 2-36

2.5.8 Electrodialysis Reversal ................................... 2-40

2.6 Alternative Technologies ......................................... 2-43

2.6.1 Iron Oxide Coated Sand .................................... 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 .......................................... 2-48

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 W/W Cost Model .................................... 3-8

3.2.4 Cost Indices and Unit Costs .................................. 3-8

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 ......................................... 3-13

3.4 Costs for Multiple Removal Percantages ............................. 3-16

3.4.1 Removal and Accessory Costs ............................... 3-16

3.4.2 Use of Blending in Cost Estimates ............................ 3-18

3.5 Pre-Oxidation Processes ......................................... 3-18

3.5.1 Potassium Permanganate .................................... 3-19

3.5.2 Chlorination ............................................. 3-20

3.6 Precipitative Processes .......................................... 3-23

3.6.1 Coagulation/Filtration ...................................... 3-23

3.6.2 Enhanced Coagulation ..................................... 3-24

3.6.3 Direct Filtration .......................................... 3-28

3.6.4 Coagulation Assisted Microfiltration .......................... 3-28

3.6.5 Lime Softening ........................................... 3-32

3.6.6 Enhanced Lime Softening ................................... 3-33

3.7 Adsorption Processes ............................................ 3-36

3.7.1 Activated Alumina ........................................ 3-36

3.7.2 Granular Ferric Hydroxide .................................. 3-47

3.8 Ion Exchange Processes .......................................... 3-47

3.8.1 Anion Exchange .......................................... 3-47
iii

3.9 Separation Processes ............................................ 3-56

3.9.1 Microfiltration ........................................... 3-56

3.9.2 Ultrafiltration ............................................ 3-56

3.9.3 Nanofiltration ............................................ 3-56

3.9.4 Reverse Osmosis ......................................... 3-56

3.10 Greensand Filtration ............................................. 3-57

3.11 Comparison of Costs ............................................ 3-61

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 ........................... 4-3

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 ........................... 4-9

4.3.4 Sanitary Landfill Disposal ................................... 4-9

4.3.5 Hazardous Waste Landfill Disposal ........................... 4-10

4.4 Residuals Characteristics ......................................... 4-11

4.4.1 Coagulation/Filtration ...................................... 4-11

4.4.2 Enhanced Coagulation ..................................... 4-13

4.4.3 Direct Filtration .......................................... 4-14

4.4.4 Coagulation Assisted Microfiltration .......................... 4-14

4.4.5 Lime Softening ........................................... 4-15

4.4.6 Enhanced Lime Softening ................................... 4-16

4.4.7 Ion Exchange ............................................. 4-18

4.4.8 Activated Alumina ........................................ 4-19

4.4.9 Nanofiltration ............................................ 4-20

4.4.10 Reverse Osmosis ......................................... 4-21

4.5 Residuals Handling and Disposal Costs .............................. 4-22
iv

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

5.2.2 pH ...................................................... 5-2

5.2.3 Co-occurrence ............................................ 5-2

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 ................................................... 5-10

5.5.1 Cost Estimates ........................................... 5-10

5.6 Activated Alumina .............................................. 5-11

5.6.1 Cost Estimates ........................................... 5-11

6.0 REFERENCES ....................................................... 6-1

APPENDIX A VERY SMALL SYSTEMS CAPITAL COST BREAKDOWN

SUMMARIES

APPENDIX B WATER MODEL CAPITAL COST BREAKDOWN SUMMARIES

APPENDIX C W/W COST MODEL CAPITAL COST BREAKDOWN SUMMARIES

APPENDIX D BASIS FOR REVISED ACTIVATED ALUMINA COSTS

APPENDIX E BASIS FOR REVISED ANION EXCHANGE COSTS
v

LIST OF TABLES

2-1 Typical IX Resins for Arsenic Removal .................................... 2-23

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 As(V) and As(III) Removal by NF Membranes .............................. 2-34

2-8 Arsenic Removal with NF at Pilot-Scale ................................... 2-35

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 ............................... 2-42

2-14 Adsorption Tests on GFH ............................................... 2-46

3-1 TDP 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 Water 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 W/W 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 Flows Used in the Cost Estimation Process ................................. 3-12

3-12 Treatment Technology Maximum Achievable Removal Percentages .............. 3-17

3-13 Assumed Percentage of Ground Water Systems with No Treatment Regeneration ... 3-38

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
vi

LIST OF FIGURES

2-1 Pressure Driven Membrane Process Classification ........................... 2-28

3-1 Pre-oxidation - 1.5 mg/L Chlorine Capital Costs ............................. 3-21

3-2 Pre-oxidation - 1.5 mg/L Chlorine O&M Costs .............................. 3-22

3-3 Enhanced Coagulation/Filtration Capital Costs .............................. 3-26

3-4 Enhanced Coagulation/Filtration O&M Cost ................................ 3-27

3-5 Coagulation Assisted Microfiltration Capital Cost ............................ 3-30

3-6 Coagulation Assisted Microfiltration O&M Cost ............................. 3-31

3-7 Enhanced Lime Softening Capital Cost ..................................... 3-34

3-8 Enhanced Lime Softening O&M Cost ...................................... 3-35

3-9 Activated Alumina Capital Costs - Natural pH .............................. 3-41

3-10 Activated Alumina O&M Costs - Natural pH of 7-8 .......................... 3-42

3-11 Activated Alumina O&M Costs - Natural pH of 8-8.3 ......................... 3-43

3-12 Activated Alumina Capital Costs - pH Adjusted to 6.0 ........................ 3-44

3-13 Activated Alumina O&M Costs - pH Adjusted to 6.0, 23,100 BV ................ 3-45

3-14 Activated Alumina O&M Costs - pH Adjusted to 6.0, 15,400 BV ................ 3-46

3-15 Bed Volumes to Arsenic Breakthrough as a Function of Sulfate Concentration ...... 3-51

3-16 Anion Exchange (< 20 mg/L SO4) Capital Cost .............................. 3-52

3-17 Anion Exchange (< 20 mg/L SO4) O&M Cost ............................... 3-53

3-18 Anion Exchange (20-50 mg/L SO4) Capital Cost ............................. 3-54

3-19 Anion Exchange (20-50 mg/L SO4) Capital Cost ............................. 3-55

3-20 Greensand Filtration Capital Cost ........................................ 3-59

3-21 Greensand Filtration O&M Cost .......................................... 3-60

4-1 Anion Exchange (< 20 mg/L SO4) Disposal Capital Costs - POTW Discharge ...... 4-23

4-2 Anion Exchange (< 20 mg/L SO4) Disposal O&M Costs - POTW Discharge ....... 4-24

4-3 Anion Exchange (20 - 50 mg/L SO4) Disposal Capital Costs - POTW Discharge .... 4-25

4-4 Anion Exchange (20 - 50 mg/L SO4) Disposal O&M Costs - POTW Discharge ..... 4-26

4-5 Mechanical Dewatering and Non-hazardous Landfill, Coagulation-Assisted

Microfiltration, Disposal Capital Costs .................................... 4-27

4-6 Mechanical Dewatering and Non-hazardous Landfill, Coagulation-Assisted

Microfiltration, Disposal O&M Costs ..................................... 4-28

4-7 Non-Mechanical Dewatering and Non-hazardous Landfill, Coagulation-Assisted

Microfiltration, Disposal Capital Costs .................................... 4-29

4-8 Non-Mechanical Dewatering and Non-hazardous Landfill, Coagulation-Assisted

Microfiltration, Disposal O&M Costs ..................................... 4-30
vii

4-9 Non-hazardous Landfill, Activated Alumina - pH 7-8, Disposal O&M Costs ....... 4-31

4-10 Non-hazardous Landfill, Activated Alumina - pH 8-8.3, Disposal O&M Costs ...... 4-32

4-11 Non-hazardous Landfill, Activated Alumina - pH 6 (23,100 BV), 

Disposal O&M Costs .................................................. 4-33

4-12 Non-hazardous Landfill, Activated Alumina - pH 6 (15,400 BV), 

Disposal O&M Costs .................................................. 4-34

4-13 POTW Discharge - 500', Greensand Filtration, Disposal Capital Costs ........... 4-35

4-14 POTW Discharge - 500', Greensand Filtration, Disposal O&M Costs ............. 4-36

5-1 POU Capital Costs, Reverse Osmosis ...................................... 5-8

5-2 POU O&M Costs, Reverse Osmosis ....................................... 5-9

5-3 POU Capital Costs, Activated Alumina .................................... 5-13

5-4 POU O&M Costs, Activated Alumina ..................................... 5-14
viii

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

DBP 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

gpm gallons per minute

GT gravity thickening

HD hazardous waste landfill disposal
ix

HDPE high-density polyethylene

ID indirect discharge

IOCS iron oxide coated sand

IX ion exchange

kgal thousand gallons

kgpd thousand gallons per day

kWh kilowatt hour

LA land application

lb/acre pounds per acre

LS lime softening

MCL Maximum Contaminant Level

MCLG Maximum Contaminant Level Goal

MD mechanical dewatering

MDL minimum detection limit

MF microfiltration

mg/kg milligram per kilogram

mg/L milligrams per liter

MGD or mgd million gallons per day

MMAA monomethyl arsenic acid

MWCO molecular weight cut-off

MWDSC Metropolitan Water District of Southern California

NF nanofiltration

NOM natural organic matter

NIPDWR National Interim Primary Drinking Water Regulation

NPDES National Pollutant Discharge Elimination System

NPDWR National Primary Drinking Water Regulation

O&M operations and maintenance

OGWDW Office of Ground Water and Drinking Water

POE point-of-entry

POTW public-owned treatment works
x

POU point-of-use

ppb parts per billion

ppm parts per million

PPI Producer Price Index (for Finished Goods)

psi pounds per square inch

psig pounds per square inch gauge

PVC polyvinyl chloride

RCRA Resource Conservation and Recovery Act

RO reverse osmosis

scf standard cubic feet

SD sanitary landfill disposal

SDWA Safe Drinking Water Act

sf square feet

SL storage lagoons

SMI Sulfur-Modified Iron

SOC synthetic organic compound

sq ft square feet

TBLL Technically Based Local Limits

TCLP Toxicity Characteristic Leaching Procedure

TDS total dissolved solids

TOC total organic carbon

TSS total suspended solids

TWG Technologies Working Group

UF ultrafiltration

UV254 ultraviolet 254

WET Whole Effluent Toxicity

wk week

yr year

Fg/L micrograms per liter
1-1

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 hyperpigmentation, 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 HAsO4

2- 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 technologies 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, specifically activated alumina (AA);

# Ion exchange (IX) processes, specifically anion exchange;

# Membrane filtration, including reverse osmosis (RO), and electrodialysis reversal

(EDR);
1-2

# Alternative treatment processes specifcally greensand filtration; and

# Point-of-use (POU) devices.

Many of these processes were evaluated to develop cost curves for the technologies.

Discussions of the following technologies are included in this document for future consideration as

viable processes for arsenic removal:

# Iron oxide coated sand;

# Nanofiltration;

# 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 Fg/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. 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, costs, implementation issues, and regulatory options before 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. The SDWA was later amended

to require the final regualation to be issued by June 22, 2001.
1-3

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. 

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 selected treatment options.

Chapter 6.0 References - Lists the literature cited in this document, as well as additional

references which may be of interest to the reader.
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2-1

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 from drinking

water. 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 of these

processes may be technically feasible, their cost may be prohibitive.

Technologies discussed in this chapter 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 (C/F) 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 floc, which more readily filter or settle under the influence of gravity.

The coagulation/filtration process has traditionally been used to remove solids from drinking

water supplies. However, the process is not restricted to the removal of particles. Coagulants render

some dissolved species [e.g., natural organic matter (NOM), inorganics, and hydrophobic synthetic

organic compounds (SOCs)] insoluble and the metal hydroxide particles produced by the addition of

metal salt coagulants (typically aluminum sulfate, ferric chloride, or ferric sulfate) can adsorb other

dissolved species. Major components of a basic coagulation/filtration facility include chemical feed

systems, mixing equipment, basins for rapid mix, flocculation, settling, filter media, sludge handling

equipment, and filter backwash facilities. Settling may not be necessary in situations where the
2-2

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 Coagulant Type

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 and 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(III) 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(III) 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(III) 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 average

concentration of arsenic in the source water was 2.1 μg/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

flocs play an important role in arsenic removal. When ferric coagulants are added, most of the ferric
2-3

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)

demonstrated 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.

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 67 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.
2-4

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.

Effect of Initial As(III)/As(V) Concentration

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 in 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(III) greater than 0.1 mg/L, neither alum nor ferric sulfate dosed at 30 mg/L could remove As(III)

to concentrations below 0.05 mg/L. In both cases, 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 μg/L, that both As(III) and As(V)

removal was independent of initial concentration. Cheng et al. (1994) showed that As(V) 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 μg/L.

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.
2-5

Optimization Hierarchy for Coagulation/Filtration Facilities

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 As(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.

Field Studies

The field operation of two coagulation filtration plants was studied by Battelle Memorial

Institute, with funding from EPA (EPA, June 2000). One plant (Plant A) uses ozonation coupled with

coagulation/filtration to treat up to 600 mgd. The other water system (Plant B) relies on coagulation,

sedimentation, and filtration, and was designed to treat a much lower daily flow (62.5 mgd). Both

plants demonstrated the ability to consistently reduce moderately high average influent arsenic

concentrations (7.5 and 19.1Fg/L) to less than 5 Fg/L (3.5 and 4.0 Fg/L) in finished water.

Furthermore, it should be noted that these plants were not using optimal coagulant and/or polymer

doses, and were not operated at the ideal pH for arsenic removal.

Based on the observation of the field operation of these systems, adsorption and

coprecipitation of As(V) with iron and aluminum flocs appears to be the principal mechanism for

arsenic removal at these plants. As part of this study, sludge samples were collected from both Plant

A and Plant B and subjected to TCLP-testing. However, based on the results of the TCLP, these

sludges would not be characterized as hazardous wastes.

Summary

Coagulation technology can successfully achieve As(V) removals greater than 90 percent. As

noted in the field study discussed above, coagulation/filtration plants have demonstrated the capacity
2-6

to reduce arsenic levels below 5 Fg/L. Moreover, if optimal operating practices are adopted, it is

anticipated that effluent levels of less than 3 Fg/L will be obtainable.

Arsenic in the pentavalant arsenate form is more readily removed than the trivalent arsenite

form. 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. The presence of sulfates significantly

decreases As(III) removal, but only 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 commonly used by facilities treating groundwater. The

oxidation process used to remove iron and manganese leads to the formation of hydroxides that

remove soluble arsenic by precipitation or adsorption reactions.

Arsenic removal during iron precipitation is fairly efficient (Edwards, 1994). Removal of 2

mg/L of iron achieved a 92.5 percent removal of As(V) from a 10 μg/L As(V) initial concentration

by adsorption alone. Even removal of 1 mg/L of iron resulted in the removal of 83 percent of influent

As(V) arsenic from a source with 22 μg/L As(V). Indeed, field studies of iron removal plants have

indicated that this treatment can feasibly reach 3 Fg/L (see discussion below). However, the removal

efficiencies achieved by iron removal are not as high or as consistent as those realized by activated

alumina (see section 2.3.1) or ion exchange (see section 2.4).

Note, however, that arsenic removal 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 μg/L

As(V) influent concentration.

Oxidation filtration technologies may be effective arsenic removal technologies. Research of

oxidation filtration technologies has primarily focused on greensand filtration. As a result, the

following discussion focuses on the effectiveness of greensand filtration as an arsenic removal

technology.
2-7

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

of the transfer of electrons and adsorption of As(V), reduced manganese (MnII) is released from the

surface.

The effectiveness of greensand filtration for arsenic removal is dependent on the influent water

quality. Subramanian 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 mg/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.

Effect of Co-occurring Inorganic Solutes

McNeill and Edwards (1995) demonstrated that a Fe/Mn facility with 400 mg/L sulfate and

5.2 μg/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.
2-8

This analysis suggests that sulfate interferes only slightly with sorption of arsenic onto ferric iron

precipitates.

Field Studies

The field operation of two iron removal plants was studied by Battelle Memorial Institute,

with funding from EPA (EPA, August 2000). One plant (Plant A) used an iron removal process

followed by zeolite softening, while the second (Plant B) relied solely on an iron removal process.

Both plants treated a similar volume of water (1.6 mgd and 1.4 mgd, respectively), however, the

source water of the former contained an average concentration of 2,284 Fg/L of iron, while the source

water of the latter was characterized by a much lower concentration (1,137 Fg/L). Since arsenic

removal is heavily dependent on the amount of iron in the influent, it is not surprising that Plant B was

not as effective in the removal of arsenic–while Plant A consistently reduced arsenic levels by 87

percent (from 20.3 Fg/L to 3.0 Fg/L), Plant B only reduced arsenic concentrations by 74 percent (from

48.5 Fg/L to 11.3 Fg/L). Based on the data collected during the study, and additional experience, the

authors of the study concluded that the removal efficiency of Plant B could be improved by the

addition of a coagulant such as ferric chloride.

As part of the same study, sludge samples were taken from Plant A and subjected to TCLPtesting. Based on the results of these tests, the wastes would not be characterized as hazardous, even

under the strict hazardous waste regulations of California. Although sludge samples were not

collected at Plant B during this study, an earlier test result would have resulted in the exceedance of

California’s limit on total arsenic for non-hazardous wastes.

2.2.3 Coagulation Assisted Microfiltration

Arsenic is removed effectively by 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 in a

manner similar 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 floc sizes can be removed (smaller amounts of coagulants are required); and

# increased total plant capacity.
2-9

Vickers et al. (1997) reported that microfiltration exhibited excellent arsenic removal

capability. This report is corroborated by pilot studies conducted by Clifford (1997), which found

that coagulation assisted microfiltration could reduce arsenic levels below 2 Fg/L in waters with a

pH of between 6 and 7, even when the influent concentration of Fe(III) is approximately 2.5 mg/L.

These studies also found that the same level of arsenic removal could be achieved by this treatment

process even if source water sulfate and silica levels were high. Further, coagulation assisted

microfiltration can reduce arsenic levels to an even greater extent at a slightly lower pH (approximately 5.5).

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 for the reduction of disinfection byproduct (DBP) precursors for surface water

systems which have sedimentation capabilities. The enhanced process involves modifications to the

existing coagulation process such as increasing the coagulant dosage, reducing the pH, or both.

Chenget 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:

# 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 μg/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.
2-10

Note that post-treatment pH adjustment may be required for corrosion control when the process

is operated at a low pH.

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 precipitation. 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 in the U.S. for reducing hardness in 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 in reducing arsenic. As(III) or As(V)

removal by LS is pH dependent. Oxidation of As(III) to As(V) prior to LS treatment will increase

removal efficiencies if As(III) 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 recalcination 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.

Effect of Initial As(V)/As(III) Concentration

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 μg/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 μg/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)
2-11

was reduced to 0.05 mg/L when the initial concentration was 0.35 mg/L or lower, while As(III) 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 μg/L of As(V) versus about 30 percent of removal in the 150 μg/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(III) 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(III) concentration 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(III) concentration of 0.24 mg/L was removed.

Effect of pH

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 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 (III), 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(III),

however, were almost identical to removals of As(V).

Effect of Type of Precipitative Solids Formed

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 softening. 
2-12

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

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 11 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 in 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 μg/L As(V) at pH 12 indicated greater

than 95 percent As(V) removal. After spiking raw water with 32 μg/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.
2-13

Field Studies

Surveys of lime softening facilities by AWWARF (AWWARF 2000) show that the technology

should be able to produce water with less than 3 Fg/L–all five of the surveyed plants that operated

at a pH of 10.2 or higher would be able to feasibly achieve these low arsenic levels. However,

although lime softening processes are quite effective for arsenic removal when operated at high pH,

lime softening facilities may not effectively remove arsenic to low levels when operated at or below

pH 10.0. For example, a field study conducted by Battelle Memorial Institute, with funding from EPA,

found that a 10 mgd lime softening plant only reduced average total arsenic concentrations of 32.0

Fg/L by about 45 percent to an average of 16.6 Fg/L in the finished water even though the treatment

train included pre-softening oxidation and post-softening filtration (EPA, June 2000). None of the

sludge samples collected as part of this study qualified as hazardous waste based on TCLP testing.

Optimization Hierarchy for Softening Facilities

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.

Summary

Softening technology can be implemented by water systems to achieve greater than 90 percent

removal of As(V). As discussed above, a survey of softening facilities by AWWARF found that those

surveyed facilities operating at a pH of 10.2 or higher could feasibly reduce arsenic concentrations

in the treated water below 3 Fg/L.
2-14

Arsenic in the pentavalent arsenate form is more readily removed by this treatment process

than the trivalent arsenite form. The optimum pH for As(V) removal by softening is approximately

10.5 and the optimum pH for removal of As(III) is approximately 11.0. As(III) removal appears to

depend on the initial concentration of the contaminant. Initial findings indicated that As(V) removal

was also a function of its concentration in the influent. However, more recent studies have shown that

As(V) removal is actually independent of its 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 ADSORPTIVE 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 Al(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

regeneration have significant effects on the removals achieved with AA. Other factors include spent

regenerant disposal, alumina disposal, and secondary water quality.
2-15

Effect of pH

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 (BV). For the same level of

As(V) removal, a column operating at pH 6.0 treated 13,391 BV and a column operating at a pH of

9.0 treated only 800 BV. 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 a pH 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).

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.
2-16

Effect of Competing Ions

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 (AWWA, 1990). As is indicated by the

general selectivity sequence shown below (Clifford and Lin, 1995), AA preferentially adsorbs

H2AsO4

-

 [As(V)] over H3AsO3 [As(III)]:

OH- > H2AsO4

-

 > Si(OH)3O- > F- > HSeO3

-

 > TOC > SO4

2- > 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 deionized water, and the presence of 720 mg/L of sulfate

resulted in more than 50 percent less arsenic removal than that achieved in deionized water.

Effect of Empty Bed Contact Time

The operation of AA beds, and in particular the EBCT, can also play a role in arsenic

removal. EBCT represents the length 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 μg/L) were achieved for more than 2,000 hours of operation.

Comparing EBCTs, the data show that adsorption increased slightly with increasing EBCT.
2-17

Regeneration

Regeneration of AA beds is usually accomplished using a strong base solution, typically

concentrated NaOH. Relatively few BV 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. (1998) 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 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 adsorption runs to decrease by 10 to 15 percent after

each regeneration.

Field Studies

Several field studies have demonstrated that arsenic may be reduced to below 3 Fg/L using

activated alumina filtration. For example, Stewart (1991) reports on the success of a small water

utility in treating for arsenic. This utility, which serves 89 households in Bow, New Hampshire,

relies on 2 wells for its source water. Each of the wells is characterized by a high level of arsenic

contamination (mean arsenic concentrations of 62 Fg/L and 57 Fg/L). However, the utility was able

to reduce arsenic levels below the current MCL of 50 Fg/L for 106 days while producing an average
2-18

treated flow of 13,730 gpd. In fact, the data shows that this system would have been able to reduce

arsenic concentrations in treated water below 5 or 10 Fg/L.

In another study (Wang, 2000), the Battelle Memorial Institute, with funding from EPA,

examined two activated alumina plants designed to treat for arsenic. The treatment trains for both of

these plants consisted of two parallel sets of two tanks in series. The media in the lead tank of each

set was replaced when it had reached its capacity for arsenic (approximately every 1.5 years). At this

time, the second tank in the series was moved to the lead or ‘roughing’ position. Thus, in all cases,

the second, or ‘polishing,’ tank contained virgin media. Challenged by arsenic concentrations ranging

from 53 to 87 Fg/L (average concentration of 62 Fg/L), and 21 to 76 Fg/L (average concentration of

49 Fg/L), both systems reduced influent arsenic concentrations by more than 92 percent on average,

consistently maintaining effluent arsenic concentrations below 5 Fg/L. Indeed, despite such high

influent concentrations, both of these two systems were able to produce finished water with an arsenic

concentration below 3 Fg/L for an extended period of time (40 and 10 weeks, respectively). Samples

of spent media were collected from one of the treatment vessels and subjected to the TCLP test. In

all cases, the spent media easily passed the TCLP test for toxicity. The results of these tests and their

implications are discussed in more detail in chapter 4 of this document.

Media Fouling

Much like ion exchange resins, 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 headloss across

the bed. If the headloss 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 BV 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

Field studies involving AA indicate that this technology can feasibly achieve arsenic removal

to 3 Fg/L (see discussion above; Stewart 1991, Wang 2000). The operational experiences which have
2-19

been developed provide important information to be considered for AA processes; these are discussed

here. 

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. Systems operating activated alumina without regeneration will also benefit from series

operation. Such an approach will provide greater utilization of the media before it is disposed.

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, posttreatment, and regeneration practices if the process is to be operated efficiently. This presents a

problem particularly for small systems. For these reasons, it is recommended that systems utilize

disposable activated alumina rather operate the process with regeneration. 

Secondary Effects

AA processes may produce changes to the effluent water quality (EPA, 1994). When pretreatment is 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

corrosion control to raise the pH would be necessary for those systems. Some systems, especially

small systems, may choose to operate the process at the natural water pH. The two full-scale plants

studied in the ORD project (Wang, 2000) were operated at natural pH values of 8 and 8.3. While this

may not yield the optimal run length, it may be sufficient for smaller systems.
2-20

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 this 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 IX 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 BV can be treated before regeneration is required. In most cases,

regeneration of the bed can be accomplished with only 1 to 5 BV of regenerant followed by 2 to 20

BV 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 regenerant and resin disposal requirements, secondary water quality effects, and design operating

parameters.

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 IX 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
2-21

applicability of IX at a particular site. Typically, strong-base anion exchange resins are used in

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. 

HCrO4

-

 > CrO4

2- > ClO4-

 > SeO4

2- > SO42- > NO3-

 > Br- > (HPO4

2-, HAsO42-, SeO32-, CO32-) > CN-

 >

NO2

-

 > Cl- > (H2PO4-, H2AsO4

-

, HCO3

-

) > OH- > CH3COO- > FThese resins have a relatively high affinity for arsenic in the arsenate form (HAsO4

2-),

however, previous studies have shown that high TDS and sulfate levels compete with arsenate and

can reduce removal efficiency (AWWA, 1990). In general, ion exchange for arsenic removal is only

applicable for low-TDS, low-sulfate source waters. Previous studies have confirmed this

generalization; the low-sulfate/low-TDS source water in a Hanford, CA study proved to be amenable

to IX 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).

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 BV to exhaustion is limited by nitrate breakthrough.

If the sulfate level is high (about 100 mg/L), however, the number of BV 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.
2-22

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, however, may produce monovalentselective 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 (following page). Data in

Table 2-1 represent BV to exhaustion using virgin IX resins. It should be noted, however, that the

capacity 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 freshly-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).
2-23

TABLE 2-1

Typical IX Resins for Arsenic Removal

Resin Type Operating Parameters BV to Exhaustion Reference

Dowex 11 2 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

4,200 Clifford and Lin

(1986)

Ionac ASB-2 2 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

4,940

Dowex SBR-1 1 Bed Volume: 1.0 cu ft

Flowrate: 1 gpm

EBCT: 7.5 min

Depth: 3.8 ft

Sulfate/As Ratio: NR

TDS/As Ratio: NR

2,800 Hathaway and

Rubel (1987)

Fox (1989)

Ionac ASB-1

 - and -

Dowex 11

1 Bed Volume: 0.014 cu ft

Flowrate: NR

EBCT: NR

Depth: NR

Sulfate/As Ratio: 420:1

TDS/As Ratio: 9200:1

C 200

C 400-500

 [projected if As(III) is

 oxidized to As(V)]

Clifford and Lin

(1985)

C A-300E

 (bench-scale)

C A-300E

 (full-scale)

2 Bed Volume: 0.0018 cu ft

Flowrate: 0.035 gpm

EBCT: NR

Depth: 1.33 ft

Sulfate/As Ratio: 300:1

TDS/As Ratio: NR

C 1,340-1,640

C 5,000-7,000

Malcolm Pirnie

(1992)

NR = Not Reported

Another approach for minimizing 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
2-24

“counter-current flow” method also minimizes regenerant requirements, i.e. volume and concentration

(EPA, 1995).

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

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,

post-treatment corrosion control, 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 BV

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 BV of regenerant are usually required to replenish the resin, depending on the
2-25

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 Ionac ASB-2) could be regenerated equivalently

using either 2 BV of 1.0 N NaCl or 5 BV of 0.25 N NaCl in “co-current flow” operation. Also, a

rinsing cycle is required after regeneration; typically only a few BV 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. Spent regenerant 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. 

Spent brine can be treated by precipitation. Clifford and Lin (1995) have shown that arsenic

levels can be substantially reduced 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 been disposed of in a landfill. The problem

with this option is that the brine stream prior to chemical precipitation could be considered hazardous

waste. As discussed in Chapter 4, brine streams with arsenic concentrations greater than 5 mg/L can

be considered hazardous waste. Since arsenic will be even more concentrated in a brine reuse

scenario, it is unlikely to be used unless the brine could be discharged to the POTW. The TDS content

of the brine may restrict that option.
2-26

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

in 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

BV 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 Field Studies

Anion exchange processes have demonstrated the capacity to consistently reduce arsenic concentrations below 3 Fg/L. The Battelle Memorial Institute, with funding from EPA, studied two ion

exchange plants located in New England. (Wang, 2000) The treatment train of one of the plants

consisted of a potassium permanganate greensand oxidizing filter followed by a mixed bed ion

exchange system. This system was regenerated every 6 days and consistently reduced influent arsenic

concentrations of 40 to 65 Fg/L to below 5 Fg/L. Indeed, this system was able to produce finished

water with arsenic concentrations below 3 Fg/L for approximately 30 weeks.

The treatment train of the second plant also incorporated pre-oxidation, consisting of a solid

oxidizing media filter followed by an anion exchange system. However, although the influent arsenic

concentrations facing this system (19 to 55 Fg/L) were not as great as those handled by the other ion

exchange system involved in this study, effluent arsenic concentrations for this plant were not

maintained at a consistently low level, ranging from below 5 Fg/L to more than 80 Fg/L. Nonetheless,

based on the breakthrough data collected during this study, the second system could feasibly achieve

finished water levels of 5 Fg/L, or even 3 Fg/L, if it was regenerated with greater frequency (i.e.,

every 3 to 4 weeks). 

This study also demonstrated that ion exchange has only a minimal impact on water pH when

used to treat for arsenic. Following treatment, the average pH of one of the two plants was identical

to the average influent pH (pH 7.5), while the average pH of the effluent from the second plant was

only slightly below that of the influent (pH 7.3 and pH 8.3, respectively). Blending a treated portion

with an untreated portion would reduce any impact on finished water pH.
2-27

2.4.12 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 must

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 Cl-/eq 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

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 potentialdriven 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
2-28

Approx. Molecular Weight

Virus

Aquatic DOC

Metal Ion

Atomic Pollen

Bacteria

Micrometers (Log Scale)

Relative

Size of

Common

Materials

0.01 0.1 1.0 10 100

100 1000 100,000 500,000

0.001

200 20,000

Desalination

Processes

Filtration

Processes

1000

Coal Dust

Colloidal Silica

Asbestos Aqueous Salt

Sugar

Pyrogen

Radius

SEPARATION SPECTRUM

ULTRAFILTRATION MEDIA FILTRATION

MICROFILTRATION

NANOFILTRATION

REVERSE OSMOSIS

Albumin Protein

Cyst

Cryptoa

sporidium

Giardia

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 Pressure Range

MF 5 - 45 psi

UF 7 - 100 psi

NF 50 - 150 psi

RO 100 - 150 psi

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.
2-29

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),

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 that can be produced from the 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.
2-30

TABLE 2-4

Typical Recovery for Membrane Processes

Membrane Process Recovery

MF to 99%

UF to 95%

NF to 85%

RO 30-85%

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 characteristics

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 Microfiltration

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
2-31

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

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 on the above discussion on particulate arsenic occurrence. Considering this, UF

alone, like MF, may not be a viable technique 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 pore 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. 
2-32

TABLE 2-5

As(V) and As(III) Removal by UF Membranes

Membrane Type MWCO Membrane

Charge

Arsenic

Species

pH Total Arsenic

Rejection (%)

Single Element

GM2540F 8,000 ( - ) V 6.9 63

GM2540F 8,000 ( - ) V 2.0 8

GM2540F 8,000 ( - ) III 7.2 <1

GM2540F 8,000 ( - ) III 10.8 53

FV2540F 10,000 None V 6.9 3

FV2540F 10,000 None III 6.8 5

Flat Sheet

GM 8,000 ( - ) V 52

FV 10,000 None V NA

PW 10,000 None V 5

NA: Not Available

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(III) rejection was high at basic

pH and negligible at neutral pH. The uncharged FV2540F membrane showed poor rejection of both

As(V) and As(III) 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(III) 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.
2-33

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 MWCO Water Type As Species As Rejection

Desal GM2540F 8,000 High DOC GW Total As 70%

Desal GM2540F 8,000

Low DOC GW Total As 30%

Finished SW

V 47%

III 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

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 UV254 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
2-34

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 high 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.

TABLE 2-7

As(V) and As(III) Removal by NF Membranes

Membrane Type MWCO Membrane

Charge

Water

Type

Species pH Total Arsenic

Rejection

(%)

Single Element

NF 45-2540 300 ( - ) DI V 6.7 85

NF 45-2540 300 ( - ) Lake V 6.9 89

NF 45-2540 300 ( - ) DI III 6.9 12

Flat Sheet

NF 45-2540 300 ( - ) DI V NA 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
2-35

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.

TABLE 2-8

Arsenic Removal with NF at Pilot-Scale

Membrane MWCO Water Type Charge As 

Species

As 

Rejection

Single Element

Accumem 400

High DOC GW ( - ) Total As 80%

Low DOC GW ( - ) Total As 60%

NF 1 NA Finished SW NA

V >95%

III 52%

NF 2 NA Finished SW NA

V >95%

III 20%

NF 3 NA Finished SW NA

V >95%

III 30%

Array

Accumem 400 High DOC GW ( - ) Total As 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
2-36

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 keep 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 Pirnie 1992). Two NF membranes were tested: (1) NF70 manufactured 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.

Another study (Chang et al, 1994) revealed that the removal efficiency dropped significantly

during pilot-scale tests where the process was operated at more realistic recoveries. When the

membrane unit was operated at a recovery of 65%, the arsenic removal efficiency dropped to 65%

and when the recovery was increased to 90%, the arsenic removal efficiency dropped down to 16%.

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.
2-37

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

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 this testing are shown in Table 2-10. 
2-38

TABLE 2-9

Summary of Arsenic Removal with RO

Location Type Operating Parameters As Removal Reference

Eugene, OR POU C 3-5 gpd

C 90% recovery

C 20-100 psi

50% Fox, 1989

Fox and Sorg, 1987

Eugene, OR POU C 3-5 gpd

C 67% recovery

C 195 psi

below MDL Fox, 1989

Fox and Sorg, 1987

Fairbanks, AL POU C low-pressure (< 100 psi) 50% Fox, 1989

Fox and Sorg, 1987

San Ysidro, NM Pilot

(hollow fiber, cellulose

acetate)

C 50% recovery

C pH adjustment to 6.3

C antiscalent addition

93-99% Clifford and Lin, 1991

San Ysidro, NM Pilot

(hollow fiber, 

polyamide)

C 50% recovery

C pH adjustment to 6.3

C antiscalent addition

99% Clifford and Lin, 1991

San Ysidro, NM POU C 10-15% recovery 91% Fox, 1989

Fox and Sorg, 1987

Tarrytown, NY Pilot

(FilmTec BW30,

Hydranautics NCM1,

Fluid Systems TCFL)

C 15 gfd

C 10% recovery

below MDL Malcolm Pirnie, 1992

Tarrytown, NY POU Not Available 86% Rogers, 1989

Charlotte Harbor,

FL

POU 

(several membrane

types)

C 1000 gpd

C 10-60% recovery

As(V) 96-99%

As(III) 46-84%

Huxstep, 1987

Cincinnati, OH POU Not Available As(III) 73% Fox and Sorg, 1987

Hudson, NH POU Not Available 40% USEPA, 1982
2-39

TABLE 2-10

 Arsenic Removal with RO at Bench-Scale

Membrane MWCO Source Water Speciation pH As Rejection

Single Element

DK2540F 180 Deionized V 6.8 96%

DK2540F 180 Lake Water V 6.9 96%

DK2540F 180 Deionized III 6.8 5%

DK2540F 180 Lake Water III 6.8 5%

Flat Sheet

DK2540F 180 Deionized V 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. 
2-40

TABLE 2-11

Arsenic Removal with RO at Pilot-Scale

Membrane Water Type Charge As

Species

As 

Rejection

Single Element

TFCL-HR

High DOC GW ( - ) Total As >90%

Low DOC GW ( - ) Total As >80%

RO 1 Finished SW NA

V >95%

III 60%

RO 2 Finished SW NA

V >95%

III 75%

RO 3 Finished SW NA

V >95%

III 68%

RO 4 Finished SW NA

V >95%

III 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
2-41

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. In 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 manufacturer.

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.

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
2-42

communities (New Mexico State University, 1979). In one of the studies conducted at Bluewater,

New Mexico, EDR brought the level of arsenic 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 approximately 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 Concentration (mg/L)

pH 7.1 (units)

TDS 810

As(total) 0.085

Fluoride 2.4

Sulfate 36

Bicarbonate 552

Chloride 142

TABLE 2-13

Raw Water Quality for Bluewater EDR Study

Parameter Concentration (mg/L)

pH (units) 7.1

TDS 908

Na 78

Sulfate 398

Silica 16

Chloride 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.
2-43

2.6 ALTERNATIVE TECHNOLOGIES

2.6.1 Iron Oxide Coated Sand 

Iron oxide coated sand (IOCS) is a rare process which has shown 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 regenerant, 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.

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.

Effect of Arsenic Oxidation State

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.

Effect of Competing Ions

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
2-44

appear to present some competition for arsenic. The addition of 4 mg/L DOC reduced As(V) sorption

by about 50 percent.

Effect of Empty Bed Contact Time

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

μg/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 μg As/mL of media.

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.

2.6.2 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 three components: (1) finelydivided 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
2-45

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.

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 μg/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.
2-46

Table 2-14

Adsorption Tests on GFH

Units Test 1 Test 2 Test 3 Test 4

Raw Water Parameters

 pH 7.8 7.8 8.2 7.6

 Arsenate Concentration μg/L 100-800 21 16 15-20

 Phosphate Concentration μg/L 0.70 0.22 0.15 0.30

 Conductivity μS/cm 780 480 200 460

 Adsorption Capacity for Arsenate g/kg 8.5 4.5 3.2 N/D

Adsorber

 Bed Height m 0.24 0.16 0.15 0.82

 Filter Rate m/h 6-10 7.6 5.7 15

 Treatment Capacity BV 34,000 37,000 32,000 85,000

 Maximum Effluent Concentration μg/L 10 10 10 7

 Arsenate Content of GFH g/kg 8.5 1.4 0.8 1.7

 Mass of Spent GFH (dry weight) g/m320.5 12 18 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.

A field study reported by Simms et al. (2000) confirms the efficacy of GFH for arsenic

removal. Over the course of this study, a 5.3 mgd GFH plant located in the United Kingdom was

found to reliably and consistently reduce average influent arsenic concentrations of 20 Fg/L to less

than 10 Fg/L for 200,000 BV (over a year of operation) at an EBCT of 3 minutes. Despite

insignificant headloss, routine backwashing was conducted on a monthly basis to maintain media

condition and to reduce the possibility of bacterial growth. The backwash was not hazardous and

could be recycled or disposed to a sanitary sewer. At the time of replacement, arsenic loading on the

media was 2.3 percent. Leachate tests conducted on the spent media found that arsenic did not leach

from the media.
2-47

The most significant weakness of this technology appears to be its cost. Currently, GFH media

costs approximately $4,000 per ton. However, if a GFH bed can be used several times longer than

an alumina bed, for example, it may prove to be the more cost effective technology. Indeed, the system

profiled in the field study presented above tested AA as well as GFH and found that GFH was

sufficiently more efficient that smaller adsorption vessels and less media could be used to achieve the

same level of arsenic removal (reducing costs). In addition, unlike AA, GFH does not require preoxidation.

A treatment for leaching arsenic from the media to enable regeneration of GFH seems feasible,

but it results in the generation of an alkaline solution with high levels of arsenate which requires

further treatment to obtain a solid waste. Thus, direct disposal of spent GFH should be favored.

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.
2-48

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 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 Fg/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 on the removals achieved and residuals characteristics, it is expected that photooxidation 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.
3-1

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 - chlorination;

# Precipitative processes, including coagulation assisted microfiltration, enhanced

coagulation/filtration, and enhanced lime softening;

# Adsorption processes, specifically activated alumina;

# Ion exchange processes, specifically anion exchange;

The next sections provide an overview of how the costs of each technology is estimated. The

sections following include a brief description of the technology, 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 models were used to develop costs: 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 W/W Cost Model (Culp/Wesner/Culp, 1994). The

models were used for the following flow ranges; linear trends are used in the transition between

models:

# 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

# W/W Cost Model - 10 to 200 mgd
3-2

All three models rely on flows to calculate capital and operation and maintenance (O&M)

costs. In addition, the Water and W/W 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). Some processes (e.g., activated alumina and

ion exchange) have slightly different ranges due to discrepancies between the models.

Activated alumina costs are not based on these models because they assume regeneration of

the media and parallel operation of columns. Activated alumina costs in this document assume no

regeneration (disposable media) and series operation of columns. Anion exchange costs are also not

based on the Water or W/W models because the models are for nitrate removal with very high sulfate

rather than arsenic removal in the sulfate ranges under consideration. The text summarizes the

approach used for activated alumina and anion exchange costs. A more detailed description can be

found in the Appendices.

3.2.2 Technology Design Panel Recommendations

Since the 1986 Safe Drinking Water Act (SDWA) reauthorization, EPA has relied mainly on

the three cost models to estimate compliance costs for drinking water regulations. Following the

reauthorization of the SDWA in 1996, EPA 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 (TDP) 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).

Total capital costs consist of three elements: process, construction, and engineering costs.

Process costs include manufactured equipment, concrete, steel, electrical and instrumentation, and

pipes and valves. Construction costs include sitework and excavation, subsurface considerations,

standby power, contingencies, and interest during construction. Engineering costs include general

contractor overhead and profit, engineering fees, and legal, fiscal, and administrative fees. Housing

costs are specifically excluded from each of these cost category designations. Housing costs are
3-3

included in the estimates presented in this chapter, but are added to the total capital cost after

application of the TDP cost approach.

To incorporate the TDP recommendations, the cost models are used to estimate total capital

costs, and then steps are taken to adjust these costs. The TDP recommended that total capital cost

estimates be based solely on process costs; therefore, in the first step, the process-related components

of the models’ estimated costs are generated using the capital cost breakdowns presented in

Appendices A through C. In the second step, construction and engineering costs are then estimated

using the factors in Table 3-1 to arrive at total capital costs that reflect the TDP recommendations. 

Table 3-1

TDP Capital Cost Factors

System Size Process Cost

Factor

(Percent of

Total)

Construction Cost

Factor 

(Percent of Total)

Engineering Cost

Factor

(Percent of Total)

Total Cost

Factor1

(Percent of Total)

Very Small 1.00 (40%) 1.00 (40%) 0.50 (20%) 2.50 (100%)

Small 1.00 (40%) 1.00 (40%) 0.50 (20%) 2.50 (100%)

Large 1.00 (30%) 1.33 (40%) 1.00 (30%) 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 for membrane

equations. Capital cost breakdowns for all technologies costed using the VSS model are presented

in Appendix A.

The Water and W/W Cost assumptions for capital cost components vary by design and average

flow. The reports Estimation of Small System Water Treatment Costs (Culp/Wesner/ Culp, 1984)

and Estimating Treatment Costs, Volume 2: Cost Curves Applicable to 1 to 200 mgd Treatment

Plants (Culp/Wesner/Culp, 1979) were used to develop capital cost breakdown summaries for the

Water and W/W 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 for each design condition. These percentages were averaged to arrive

at a universal capital cost breakdown which could be applied for developing the capital costs. Tables

3-3 through 3-6 demonstrate the methodology described here.
3-4

Table 3-2

VSS Capital Cost Breakdown for Membrane Processes 

(Including Microfiltration and Ultrafiltration)

Cost Component Model

Assumption

Cost Factor Percent of

Total Capital

Capital Cost

Category

Manufactured Equipment 100% 1.000 56.97% p

Installation 25% 0.2500 14.24% c

Sitework and Interface Piping 6% 0.0750 4.27% c

Standby Power 5% 0.0625 3.56% c

General Contractor Overhead & Profit 12% 0.1665 9.49% e

Legal, Fiscal and Administrative Fees 3% 0.0416 2.37% e

Engineering 10% 0.1596 9.09% e

Miscellaneous and Contingencies 0% 0.000 0.00% c

TOTAL 1.7552 100.00%

p = process, c = construction, e = engineering

Output from the Water and W/W 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 W/W Cost models, respectively.

Capital cost breakdowns for each technology and unit process are presented in Appendices A through

C for the VSS, Water, and W/W Cost models, respectively.
3-5

Table 3-3

Water Model Capital Cost Breakdown for

Package Conventional Treatment (Coagulation/Filtration)

Cost Component

Filter Area (ft2) Capital

Cost

2 12 20 40 112 150 Category

Excavation and Sitework $3,500 $3,500 $4,700 $5,800 $7,000 $9,300 c

Manufactured Equipment $31,000 $44,900 $53,500 $111,300 $176,600 $190,500 p

Concrete $1,000 $1,000 $1,500 $4,500 $5,700 $6,800 p

Labor $9,900 $14,700 $17,500 $36,400 $57,800 $62,400 c

Pipes and Valves $4,200 $8,300 $10,400 $20,900 $29,200 $41,700 p

Electrical $3,200 $4,500 $5,300 $11,100 $17,600 $19,000 p

Housing* $18,600 $18,600 $23,400 $45,000 $47,500 $52,500 c

Subtotal $71,400 $95,500 $116,300 $235,000 $341,400 $382,200

Contingencies $10,700 $14,300 $17,400 $35,300 $51,200 $57,300 e

Total $82,100 $109,800 $133,700 $270,300 $392,600 $439,500

*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

Filter Area (ft2) Average

2 12 20 40 112 150 Percent

Excavation and Sitework 4.26% 3.19% 3.52% 2.15% 1.78% 2.12% 2.84%

Manufactured Equipment 37.76% 40.89% 40.01% 41.18% 44.98% 43.34% 41.36%

Concrete 1.22% 0.91% 1.12% 1.66% 1.45% 1.55% 1.32%

Labor 12.06% 13.39% 13.09% 13.47% 14.72% 14.20% 13.49%

Pipes and Valves 5.12% 7.56% 7.78% 7.73% 7.44% 9.49% 7.52%

Electrical 3.90% 4.10% 3.96% 4.11% 4.48% 4.32% 4.15%

Housing* 22.66% 16.94% 17.50% 16.65% 12.10% 11.95% 16.30%

Contingencies 13.03% 13.02% 13.01% 13.06% 13.04% 13.04% 13.03%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

*Housing costs are added to the total capital cost after application of the TDP cost approach
3-6

Table 3-5

 W/W Cost Model Capital Cost Breakdown for Sedimentation Basins

Cost Component

Area (A = ft2) and Length x Width (LW = ft x ft)

Capital

Cost

Category

A=240

LW = 30x8

A=600

LW=60x10

A=1260

LW=90x14

A=2240

LW=140x1

6

A=3600

LW=200x1

8

A=4800

LW=240x2

0

Excavation and Sitework $1,060 $2,000 $3,060 $4,680 $6,670 $8,090 c

Manufactured Equipment $8,540 $12,080 $24,470 $32,020 $53,110 $63,440 p

Concrete $2,970 $5,490 $84,430 $12,820 $19,190 $22,070 p

Steel $6,400 $13,110 $19,440 $32,620 $51,250 $39,680 p

Labor $6,220 $11,260 $17,320 $26,390 $37,570 $45,300 c

Pipes and Valves $6,960 $7,400 $9,100 $12,500 $16,100 $21,450 p

Electrical $1,510 $1,760 $1,860 $2,020 $2,110 $2,400 p

Subtotal $33,660 $53,100 $83,680 $123,050 $190,000 $232,430

Contingencies $5,050 $7,970 $12,550 $18,460 $27,750 $34,860 e

Total $38,710 $61,070 $96,230 $141,510 $212,750 $267,290

Table 3-6

 W/W Cost Model Capital Cost Breakdown by Percentage for Sedimentation Basins

Cost Component

Area (A = ft2) and Length x Width (LW = ft x ft)

Average

Percent

A=240

LW =

30x8

A=600

LW=60x10

A=1260

LW=90x14

A=2240

LW=140x1

6

A=3600

LW=200x1

8

A=4800

LW=240x2

0

Excavation and Sitework 2.74% 3.27% 3.18% 3.31% 3.14% 3.03% 3.11%

Manufactured Equipment 22.06% 19.78% 25.43% 22.63% 27.96% 23.73% 23.10%

Concrete 7.67% 8.99% 8.76% 9.06% 8.55% 8.26% 8.55%

Steel 16.53% 21.47% 20.20% 23.05% 24.09% 26.07% 21.90%

Labor 16.07% 18.44% 18.00% 18.65% 17.66% 16.95% 17.63%

Pipes and Valves 17.98% 12.12% 9.46% 8.83% 7.57% 8.02% 10.66%

Electrical 3.90% 2.88% 1.93% 1.43% 0.99% 0.90% 2.01%

Contingencies 13.05% 13.05% 13.04% 13.05% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%
3-7

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

DES = 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.
3-8

3.2.3.2 Water Model

1. 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.

$376,138 x 2.5 = $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 percent of the construction cost. The total capital cost is:

($692,066 x 0.1630) + $940,345 = $1,053,152

3.2.3.3 W/W Cost Model

1. The W/W 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 then be calculated using the total cost factor presented in

Table 3-1.

$275,897 x 3.33 = $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.

3.2.4 Cost Indices and Unit Costs

Both the Water Model and the W/W Cost Model require a number of standard indices and

various unit costs from the Bureau of Labor Statistics, the Engineering News Record, and other
3-9

sources. The September 1998 index values used to develop cost estimates are reported in Tables 3-7

through Table 3-9.

Table 3-7

Costs Indices Used in the Water and W/W Cost Models

Description Index

Reference

Numerical

Value1

Concrete Ingredients and Related Products BLS 132 448.8

Electrical Machinery and Products BLS 117 281.8

General Purpose Machinery and Equipment BLS 114 445.1

Metals and Metal Products (Steel) BLS 1017 405.1

Miscellaneous General Purpose Equipment

(Pipes &Valves)

BLS 1149 521.5

PPI Finished Goods Index BLS 3000 364.0

ENR Building Cost Index 3375.31

ENR Skilled Labor 5317.36

ENR Materials Prices 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

Electricity1$0.08/kWh

Diesel Fuel1$1.25/gallon

Natural Gas1$0.006/scf

Labor2 Large systems: $40/hr

Small systems: $28/hr

Building Energy Use 102.6 kWh/ft2/yr

1

 Energy Information Administration.

`
3-10

TABLE 3-9

Chemical Costs

Chemical Cost Units

Alum, Dry Stock $300 per ton

Carbon Dioxide, Liquid $340 per ton

Chlorine, 1 ton cylinder $350 per ton

Chlorine, 150 lb cylinder $400 per ton

Chlorine, Bulk $280 per ton

Ferric Chloride $350 per ton

Hexametaphosphate $1276 per ton

Lime, Quick Lime $95 per ton

Phosphoric Acid $300 per ton

Polymer $2.25 per lb

Potassium Permanganate $2700 per ton

Soda Ash $400 per ton

Sodium Hypochlorite, 12% $1100 per ton

Sodium Chloride $99 per ton

Sodium Hydroxide, 50% solution $371 per ton

Sulfuric Acid $116 per ton

The models present total capital costs and annual O&M costs. Annual O&M costs include the

costs for materials, chemicals, power, and labor. Annual costs can be determined using the following

equations:

Total annual cost (¢/kgal) = Annualized Capital Cost (¢/kgal) + O&M Cost (¢/kgal)

Where:

Annualized Capital Cost = Capital Cost ($) * Amortization Factor * 100 ¢/$

Average Daily Flow (mgd)*(1000 kgal/mgal)*365 days/year

O&M Cost (¢/kgal) = Annual O&M ($) * 100 (¢/$)

Average Daily Flow (mgd)*1000 kgal/mgal*365 days/year
3-11

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 = (1 + i)N

((1 + i)N - 1)/i)

Where: i = interest rate

N = number of years

Table 3-10 

Amortization Factors

Interest Rate (%) Amortization Period Amortization

Factor

3 20 0.0672157

7 20 0.0943929

10 20 0.1174596

3.2.5 Re-Basing Bureau of Labor Statistics Cost Indices

The Water Model and W/W Cost Model uses BLS cost index information based in 1967

dollars. In 1986, the BLS conducted a comprehensive overhaul of the industrial price methodology

resulting in a re-basing of all index information to a base year of 1982 (i.e., 1982=100). The BLS

index information must be re-based to 1967 to use the indices in the models. A sample re-base

calculation is presented below.

Sample Re-base Calculation:

Machinery = 1982 Base Factor / Re-base Factor = 1967 Base Factor 

= 146.4 / 0.32895016 = 445.1
3-12

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 flows were used for each of the

flow regions, with the exception of the transition regions, where cost estimates are based on a linear

regressions between the last data point of the previous region and the first data point of the following

region. Table 3-11 presents the design and average flows, and cost models used in this process.

Each model is used to estimate the cost for each flow. Regression analysis is then used to estimate

an equation relating cost to flow for each region.

 

Table 3-11

Flows Used in the Cost Estimation Process

Design Flow (mgd) Average Flow (mgd) Cost Model

0.010 0.0031 VSS

0.024 0.0056 VSS

0.087 0.024 VSS

0.10 0.031 VSS

0.27 0.086 Water

0.45 0.14 Water

0.65 0.23 Water

0.83 0.30 Water

1.0 0.36 Water

1.8 0.7 W/W Cost

4.8 2.1 W/W Cost

10 4.5 W/W Cost

11 5 W/W Cost

18 8.8 W/W Cost

26 13 W/W Cost

51 27 W/W Cost

210 120 W/W Cost

430 270 W/W 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, Cryptosporidium), 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. Selection of treatments that reduce additional contaminants would

generate additional benefits from the treatment technology. 

Land

Land costs are not included as part of the construction and capital costs for installation of new

facilities. Land costs and the amount of land needed will vary significantly from site to site.

Technology selection and the footprint of the technology will also play a critical role in land needed

for treatment. The November 1999 Technology and Cost Document presented ranges of costs based

on the Technology Design Panel recommendations. Comments on the proposed rule provided very

little documentation on land costs for compliance with arsenic. One comment from the Association
3-14

of California Water Agencies indicated that land costs could be as high as 5% of total annualized

costs. This is based on land prices in California, so it represents an upper bound for the nation.

It is expected that systems with existing treatment will usually have adequate land available.

Furthermore, small ground water systems are not expected to have land acquisition requirements.

Intermediate to large ground water systems and large surface water systems with appreciable ground

water use may need to acquire land. However, according to SDWIS, these systems comprise only

about five percent of all community water systems in the United States. 

Based on the above discussion, 5% of the capital costs was selected as the upper bound

increase associated with land at ground water sites. This upper bound can be used in a sensitivity

analysis to evaluate an upper bound impact of land on costs. Many systems may not need to purchase

land for arsenic treatment. 

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. 

Permitting

The Very Small Systems Document, the Water Model and the W/W Model all include costs

for legal, fiscal and administrative costs. These costs vary by model and system size. Legal, fiscal

and administrative costs are included in the engineering cost factor. For the base case, permitting

costs are covered by the legal, fiscal and administrative costs in the models. 

An upper bound for sensitivity analysis was based on the Technology Design Panel

recommendation of permitting costs as 3% of construction value (EPA, 1998). Participants in the TDP

also recommended a floor of $2500 for systems serving less than 10,000 people. The upper bound that

can be used to assess sensitivity of the costs to permitting is an additional 3% of the capital costs

rather than 3% of the construction value. This approach will yield a higher cost for permitting.
3-15

Piloting

The Technology Design Panel stated that pilot testing needed to be part of EPA’s high cost

models or upper bound estimates (EPA, 1998). Members of the large system breakout group indicated

that large systems will need to conduct pilot tests in all cases for treatment optimization. The costs

in this document assume that pilot testing is covered by the existing capital costs due to the

conservative assumptions used to develop these costs. The technology costs presented in this chapter

do not assume optimized design and operation at each site. Systems which incur piloting costs will

generally achieve more optimized, lower cost, technology applications.

An upper bound for piloting was also estimated. In other cost analyses, EPA has assumed that

piloting costs would be approximately 3% of the total capital costs for a technology. The sensitivity

analysis can use 3% of the capital costs as an upper bound.

Waste Disposal

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, 1993a) and Water System Byproducts Treatment and Disposal Cost Document

(DPRA, 1993b). 

Storage/Building

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

There are redundant items included in the cost equations for several of the technologies in

Chapter 3. For example, redundant columns were included in the costs for anion exchange and

activated alumina. Thus, redundancy is included for the major cost component in the design for these
3-16

two technologies. Redundancy for the lower cost components would be covered by the conservative

assumptions used in deriving the capital costs.

Recommended Standards for Water Works (Great Lakes Upper Mississippi River Board of

State Public Health and Environmental Managers, 1997), often referred to as the Ten State Standards,

presents a comprehensive discussion of redundancy and recommended redundant items for water

treatment facilities in those States. Systems may incur additional costs if additional redundant items

are required by the State.

3.4 COSTS FOR MULTIPLE REMOVAL PERCENTAGES

Capital and O&M cost estimates are presented for the maximum achievable removal in this

document. Table 3-12 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

Both the April 1999 and the November 1999 drafts of the Technology and Cost Document

included costs for accessories in Appendix D. Accessory costs included 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. Inclusion

of accessory costs was not based on a recommendation of the Technical Design Panel. A detailed

examination of Section 3.11 of the November 1999 Technology and Cost Document revealed the

apparent source of the accessory costs. Section 3.11 compared the costs in the November 1999

Technology and Cost Document with the costs from the 1993 Technology and Cost Document

(Malcolm Pirnie, 1993a) and Evaluation of Central Treatment Options as Small System Compliance

Technologies (SAIC, 1999). There is a statement in bold at the end of Section 3.11.1 that states that

the 1993 estimates are greater than the 1999 estimates for many technologies. It further states that this

is because the 1993 estimates included accessory costs; i.e., raw and finished water pumping and

clearwell storage.
3-17

The 1993 Technology and Cost Document was examined to determine why the accessory costs

were included for each technology. After examining the costs in this document, accessory costs, such

as clearwell storage, were only included for one technology - coagulation/filtration. This is

understandable because coagulation/filtration would not be installed solely for arsenic removal. This

technology would likely be selected by surface water systems, especially since the surface water

treatment rule (SWTR) had only promulgated only recently before the 1993 Technology and Cost

Document was prepared. Surface water systems would need the clearwell storage for disinfection

credit under the SWTR. The costs for lime softening in the 1993 Technology and Cost Document

specifically state that clearwell storage is not included. Costs were also not included for anion

exchange or activated alumina. Therefore, cost estimates presented in this document do not

include accessory capital and O&M.

Table 3-12

Treatment Technology Maximum Achievable Removal Percentages

Treatment Technology Maximum Percent

Removal

Coagulation/Filtration 95

Enhanced Coagulation/Filtration195

Coagulation Assisted Microfiltration 90

Lime Softening (pH > 10.5) 90

Enhanced Lime Softening1 (pH > 10.5) 90

Ion Exchange (sulfate < 50 mg/L) 95

Activated Alumina 95

Reverse Osmosis >95

Greensand Filtration (20:1 iron:arsenic) 80

POU Activated Alumina 90

POU Ion Exchange 90

1 - Enhanced processes assume the existing plant can achieve 50 percent removal without modification. Process

enhancements result in the balance to achieve the maximum removal. For example, an existing

coagulation/filtration facility can achieve 50 percent removal. Process enhancements result in an additional 45

percent removal, for a total removal of 95 percent.
3-18

3.4.2 Use of Blending in Cost Estimates

Capital and O&M costs were estimated using the VSS, Water, and W/W 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 percent removal. If the raw water arsenic concentration is 20 Fg/L

and the treatment objective is a finished water concentration of 10 Fg/L, the utility need only remove

50 percent of the arsenic in the raw water. In 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:

Qtreated = Qtotal 

 [((Cmax - Cdesired)/Cdesired) + 1]

Where: Qtreated = Treated portion of the total process flow, mgd

Qtotal = Total daily process flow, mgd

Cmax = Maximum achievable removal efficiency, %

Cdesired = Desired removal efficiency, %

If 1 is substituted for the total daily flow (Qtotal) 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 As(V), 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 on 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
3-19

cause disinfection by-product (DBP) 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 permanganate is not known to form

measurable DBPs and does not foul membranes. Raw water and downstream process considerations

should be made when selecting a pre-oxidation technology. Potassium permanganate costs can be

calculated for dosages of 10 mg/L. The Very Small Systems Best Available Technology Cost

Document (Malcolm Pirnie, 1993) can be used for calculating costs for the flows below 1 mgd. For

flows greater than 10 mgd the W/W Cost Model can be used to estimate the capital and O&M costs.

Linear regressions can be used to estimate costs in the transition regions between the two models, i.e.,

1 to 10 mgd. The following are some of the system design criteria that will affect costs when using

the two models:

# 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 percent potassium permanganate solution.

# The VSS document makes provisions for building (42.7 percent), fencing (49.4 percent),

and road (33.8 percent) costs associated with potassium permanganate addition. It should

be assumed that permanganate addition will be installed as part of a larger treatment

process and that building, fence and road costs for the treatment facility would be adequate

to accommodate permanganate addition.

# O&M costs for very small systems can be calculated using equations in the very small

systems cost document. Labor requirements should assume 1 hour per week of

incremental labor.

# For small system potassium permanganate addition, a dry chemical feed system capable

of 1,000 pounds per day can be used.

The costs associated with potassium permanganate are in the costs of each of the treatment

technologies considered below. 
3-20

3.5.2 Chlorination

As previously stated, chlorination can cause DBP 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 mg/L. Similar to potassium

permanganate systems, the VSS Model was used for calculating costs for the flows below 1 mgd, and

the W/W 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 percent sodium hypochlorite feed stock.

# The VSS Model makes provisions for building (52.2 percent), fencing (60.5 percent), and

road (41.4 percent) costs associated with chlorine addition. Housing costs are only

needed when the arsenic removal process is a centrally-managed point-of-use option.

Otherwise, it is assumed that chlorination process would be housed in the same building

as the larger central treatment process.

# Capital costs were calculated for both with and without housing costs added.

# Incremental 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 3-1 and 3-2.
3-21

y = 7277

y = 6875x + 401.97

y = 15290x0.6554

y = 14560

y = 6065.8x + 8494.2

$1,000

$10,000

$100,000

$1,000,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

With Housing

Without Housing

Figure 3-1

Pre-oxidation - 1.5 mg/L Chlorine

Capital Costs
3-22

y = -1E-10x2 + 4239.7x + 1161.8

y = 4920.1x + 916.87

y = -0.3928x2 + 823.35x + 19360

$100

$1,000

$10,000

$100,000

$1,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-2

Pre-oxidation - 1.5 mg/L Chlorine

O&M Costs
3-23

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. C/F is unlikely to

be installed solely for arsenic removal. Systems considering installation of this technology should

design the process to operate in the optimal pH range if high removal efficiency is needed for

compliance. Source waters containing high As(III) concentrations may opt for oxidation as a pretreatment for C/F. Pre-oxidation options are presented in Section 3.5. Costs are not presented for this

option. Key design parameters that can be used with the cost models are presented for systems that

wish to explore this option.

Very Small Systems (Less than 0.10 mgd)

Capital costs for very small systems can be 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 percent), fencing (2.1-7.5 percent) and road (1.2-4.7 percent) costs

associated with each of the technologies presented. The following design criteria can be 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 can be escalated using the BLS

Chemical and Allied Products Index. Labor requirements were estimated at 8 hours per week.
3-24

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 can be used in developing capital and O&M cost estimates:

# Package plant for all small systems, filtration rate 5 gpm/ft2;

# 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 W/W Cost model can be used to develop capital and O&M cost estimates for large C/F

plants. The following design criteria can be 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.

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 preoxidation 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 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.

Small Systems (Less than 1 mgd)

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
3-25

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 3-3 and 3-4 present capital and O&M cost curves and equations for enhanced

coagulation. 
3-26

y = -3825.4x2 + 14734x + 7144.5

y = 94000x - 75947

y = -6.7115x2 + 6482.1x + 799902

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

Figure 3-3

Enhanced Coagulation/Filtration

Capital Costs
3-27

y = -302.31x2 + 7299x + 220.79

y = 17479x - 3483.3 y = -0.3972x2

 + 14950x + 7906.4

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-4

Enhanced Coagulation/Filtration

O&M Costs
3-28

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. Preoxidation technologies are discussed in Section 3.5. Direct filtration is unlikely to be installed solely

for arsenic removal. The technology may not be able to achieve high removals of arsenic. Design

criteria are not presented for direct filtration.

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 floc 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.

Very Small Systems (Less than 0.10 mgd)

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 given 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 on vendor

quotes and case studies. These costs were then added to the C/F cost estimates.
3-29

Small Systems (Less than 1 mgd)

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.

Large Systems (Greater Than 1 mgd)

The W/W 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 3-5 and 3-6 present capital and O&M cost curves and equations for removal of arsenic

by coagulation assisted microfiltration.
3-30

y = -11935465x2 + 4880036x + 94324

y = 2343199x + 228653

y = -483591x2 + 2308991x + 273143

y = 1030810x + 1067733

y = 320x2 + 921471x + 2129119

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

$1,000,000,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

Figure 3-5

Coagulation Assisted Microfiltration

Capital Costs
3-31

y = 196829x + 20264

y = 136332x + 22139

y = 80081x + 26977

y = 13311x + 51014

y = 15236x + 42350

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-6

Coagulation Assisted Microfiltration

O&M Costs
3-32

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.

Lime softening is unlikely to be installed solely for arsenic removal. Systems considering

installation of this technology should design the process to operate in at a pH > 10.5 if high removal

efficiency is needed for compliance. Costs are not presented for this option. Key design parameters

that can be used with the cost models are presented for systems that wish to explore this option.

Considerable amounts of sludge are produced by the LS process. Large systems may find it

economically feasible to install recalcination equipment to recover and reuse the process sludge and

reduce disposal costs. 

Small Systems (Less than 1 mgd)

The VSS Model provides no estimation methods for LS treatment. Therefore, the Water Model

would have to be used to estimate capital and O&M costs for all LS treatment plants with less than

1 mgd capacity. The following design criteria can be 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.

Large Systems (Greater Than 1 mgd)

The W/W Cost model can be used to develop capital and O&M cost estimates for large LS

plants. The following design criteria can be 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.
3-33

3.6.6 Enhanced Lime Softening

Enhanced LS involves modifications to the typical LS treatment process to adjust the pH >

10.5 by 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.

Small Systems (Less than 1 mgd)

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 3-7 and 3-8 present cost curves and equations for removal of arsenic by enhanced lime

softening.
3-34

y = -22974x2 + 48211x + 7897.7

y = 133486x - 100351

y = -7.539x2 + 26610x + 969164

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

Figure 3-7

Enhanced Lime Softening

Capital Costs
3-35

y = 2144.7x2 + 30525x + 319.67

y = 29145x + 1094.3 y = -0.4528x2 + 26067x + 14956

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-8

Enhanced Lime Softening

O&M Costs
3-36

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 passed through the bed to remove

contaminants. The contaminant ions are exchanged with the surface hydroxides on the alumina.

Disposable activated alumina (no regeneration) costs were developed using the following approach.

Four sets of costs were developed - two for unadjusted pH and two where the pH has been adjusted

to the optimal pH of 6. The costs for pH adjustment to optimal pH 6.0 are included as a separate

module in these cost estimates since it is only used for two of the options. 

The three models routinely used to estimate unit costs are all based on regeneration of the

media, operation at optimal pH, and use of single columns. Most of the costs are based on fluoride

removal rather than arsenic removal. It was determined that the existing cost models could not be used

to estimate costs for disposable activated alumina. To estimate costs, it was assumed that the alumina

would be arrayed in small columns operated in series rather than in one larger column. This will

provide greater utilization of the media before disposal and is more consistent with the other designs

using disposable activated alumina. 

Design Assumptions: Capital Costs with Natural pH

The following assumptions were made to estimate the capital cost of activated alumina when

the water is not adjusted for pH. The basis for these assumptions are provided in Appendix D. 

 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 ranges will be

evaluated 7.0 #pH < pH 8.0 and 8.0 # pH #8.3.

 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.
3-37

 3. Empty Bed Contact Time (EBCT) is 5 minutes per column

 4. The estimated cost of the activated alumina media is $0.82/lb for systems with a design flow

greater than 0.1 mgd. The estimated cost of activated alumina media is $1.30/lb for systems

with a design flow less than or equal to 0.1 mgd.

 5. The density of the activated alumina media is assumed to be 47 lb/ft3.

 6. The bed depth assumption ranged from 3 to 6 feet, depending on the design flow. The

maximum diameter per column was 12 feet. When the calculated media volume exceeded the

maximum, the treatment media was divided into several beds to maintain a realistic bed

diameter. Thus, larger designs include multiple activated alumina trains, each treating a

portion of the total flow.

 7. The vessel cost has been sized based on 50 percent bed expansion during backwash even

though backwashing may not be necessary on a routine basis for smaller systems. The vessel

volume was calculated as (1.5)(media volume). The vessel cost is given by the following

equation: 

Cost = 63.288 * (Vessel Volume in gallons) ^ 0.679

 8. The capital costs include a redundant column to allow the system to operate while the media

is being replaced in the old roughing column.

 9. The remaining components of the process equipment include concrete, pipe and valves and

electrical and instrumentation. A cost factor of 14.43 percent of the manufactured equipment

and activated alumina was used for these other construction cost components.

 10. The capital costs have been estimated from the total process costs using a factor of 2.5 for

small systems and 3.33 for large systems.
3-38

 11. 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.

Table 3-13 shows the assumptions made for this analysis regarding the percentage of systems

with no treatment in place. (See Appendix D for further information on these percentages.)

Table 3-13

Assumed Percentage of Ground Water Systems with No Treatment

Population Category Percentage of GWS w/o any Treatment

25 - 100 43%

101 - 500 19%

501 - 1,000 16%

1,001 - 3,300 18%

3,301 - 10,000 18%

10,001 - 50,000 18%

50,001 - 100,000 18%

100,001 - 1 MIL 18%

 12. 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 = 726.6 * Design Flow + 59

 13. The building cost to house the activated alumina system is $45.38/sq ft for design flows at or

below 0.65 mgd. The following equation is used for design flows between 0.65 and 1 mgd:

Building cost = 54.204 * Process Area + 14887
3-39

For design flows greater than 1 mgd, the building cost varied based on flow from $100 to $90

per square foot.

Design Assumptions: Capital Costs with pH Adjustment

For two options, the water pH is adjusted to the optimal pH of 6. The capital cost equation

for the pH adjustment module, which includes post-treatment corrosion control is given by the

following equation (Appendix D provides additional information): 

 Capital cost = 40517 * Design Flow + 31889

Design Assumptions: O&M Costs with and without pH Adjustment

There are three major components to the operation and maintenance (O&M) costs -

replacement media, energy and labor.

 1. 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. The run

length for the natural pH range of 7 # pH < 8 is 10,000 bed volumes (BV). The run length for

the natural pH range of 8 # pH # 8.3 is 5,200 BV. These run lengths are used to estimate the

replacement of the activated alumina in half of the active columns. 

 2. The O&M costs were calculated at 2 run lengths when pH is adjusted to pH of 6. The two run

lengths that were selected for the optimal pH of 6 are 15,400 and 23,100 BV. 

 3. Chemical costs also are incurred to adjust the pH. The following equation was used for

chemical costs for pH adjustment:

Chemical cost (pH adjustment) = 81361* Average Flow + 151.7

 4. The second major component in the O&M costs is incremental labor cost. Incremental labor

is the labor associated with the additional maintenance that comes with a new process.
3-40

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 5 percent breakthrough) depending

on the volume of media being replaced. It was assumed that sixteen hours are required to

replace the media for design flows less than 1 mgd and thirty-two hours are required for larger

design flows. Incremental labor for the pH adjustment processes also is included in the two

options operated at optimal pH. The incremental labor is an additional two hours per week.

 5. The labor rate for small systems was $28/hour. The loaded labor rate for large systems was

$52/hour.

 6. 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 on an annual usage of 19.5 kwh/sq 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 $0.08/kwh.

The capital and O&M cost curves for activated alumina are shown in figures 3-9 through 3-14.
3-41

y = 515309x + 10214

y = 420286x + 10212

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

$1,000,000,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

Figure 3-9

Activated Alumina

Capital Costs
3-42

y = 188890x + 4122.7

y = 190726x + 9798.2

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-10

Activated Alumina (pH 7 - 8)

O&M Costs
3-43

y = 359695x + 6388.6

y = 364399x + 15765

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-11

Activated Alumina (pH 8 - 8.3)

O&M Costs
3-44

y = 555826x + 42103

y = 460803x + 42101

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

$1,000,000,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

Figure 3-12

Activated Alumina (pH Adjusted to 6.0)

Capital Costs
3-45

y = 165316x + 5794.4

y = 165389x + 11692

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-13

Activated Alumina (pH Adjusted to 6.0 - 23,100 BV)

O&M Costs
3-46

y = 205368x + 6325.7

y = 206113x + 13092

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-14

Activated Alumina (pH Adjusted to 6.0 - 15,400 BV)

O&M Costs
3-47

3.7.2 Granular Ferric Hydroxide

Granular ferric hydroxide is a technology that may allow for very long run length without the

need for pH adjustment or pre-oxidation. It has been demonstrated for full-scale arsenic removal in

England (Simms, et. al., 2000). It was not designated a BAT because of the lack of published data

showing performance for a range of water qualities. Also, there is little published cost data, so costs

are difficult to evaluate and are not presented. 

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

(IX 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 multimedia 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 developed in this section. It is further

assumed that corrosion control measures are not required because the pH is not significantly affected

by the process, especially when blending can be utilized.

Sulfate concentrations in the influent water significantly affect the capacity of the IX resin with

respect to the removal of arsenic. Clifford (1995) estimated bed volumes for 10 percent and 50

percent breakthrough of influent arsenic as a function of influent sulfate concentration. Figure 3-15

shows the bed volume and sulfate relationship estimated by Clifford (1995). Using this figure, the

regeneration frequency for an IX column can be estimated for the entire range of sulfate concentration

based on the upper bound concentration. This is a very conservative assumption and underestimates

run length for systems with lower sulfate concentrations in the range. The run length until regeneration

is the critical factor in estimating the O&M cost. Capital costs will be marginally affected by changes

in bed volume to regeneration.
3-48

The costs in both the April 1999 Technology and Cost Document (1) and the November 1999

Technology and Cost Document (2) utilize models based on nitrate removal. The Water Model was

one of the cost models used for small systems with design flows less than 1 mgd. Sulfate

concentrations for this model were between 50 and 100 mg/L. The November 1999 Technology and

Cost Document listed the design criteria for large systems (> 1 mgd) using the W/W model. The

sulfate concentration was 80 mg/L and the nitrate concentration was 100 mg/L. It was not possible

to modify the sulfate concentration in either of these two models. It was determined that the existing

cost models could not be used to estimate costs for anion exchange options with sulfate at or below

50 mg/L. 

Design Assumptions: Capital Costs

The following assumptions were made to estimate the capital cost of anion exchange. The

basis for these assumptions are provided in Appendix E. 

 1. Empty Bed Contact Time (EBCT) = 2.5 minutes per column

 2. The estimated cost of the anion exchange resin was $125/ft3 for all system sizes.

 3. The bed depth assumption ranged from 3 feet to 6 feet depending upon the design flow. The

maximum diameter per column was 12 feet. When the calculated media volume exceeded the

maximum, the treatment media was divided into several beds to maintain a realistic bed

diameter. Thus, larger designs include multiple anion exchange columns, each treating a

portion of the total flow.

 4. The vessel cost has been sized based on 50% bed expansion during backwash. The vessel

volume was calculated as (1.5)(media volume).

 5. The vessel cost is based on the following equation: 

Cost = 63.288 * (Vessel Volume in gallons) ^ 0.679
3-49

 6. The capital costs include a redundant column to allow the system to operate while the media

is being regenerated in the other column.

 7. The other equipment cost elements are the brine tank, brine pump and mixer. Costs for the

brine tank were based on 30-day storage. Separate costs were prepared for the two sulfate

ranges based on the number of regenerations per month. 

 8. The remaining components of the process equipment include concrete, pipe and valves and

electrical and instrumentation. A cost factor of 25.02% of the manufactured equipment and

activated alumina was used for these other construction cost components.

 9. The capital costs have been estimated from the total process costs using a factor of 2.5 for

small systems and 3.33 for large systems.

 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.

Table 3-13 (above) shows the assumptions made for this analysis regarding the percentage of

systems with no treatment in place. 

 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 = 118.7 * Design Flow + 118

 12. The building cost to house the anion exchange system is $45.38/sq ft for design flows at or

below 0.65 mgd. The following equation is used for design flows between 0.65 and 1 mgd:

Building cost = 54.204 * Process Area + 14887
3-50

For design flows greater than 1 mgd, the building cost varied based on flow from $100 to $90

per square foot.

Design Assumptions: Operation and Maintenance (O&M) Costs

There are major components to the operation and maintenance (O&M) costs are regeneration

frequency, regeneration dose, and incremental labor. See Appendix E for more detail.

 1. The critical variable in calculating the regeneration frequency is the run length - the number

of bed volumes until the effluent concentration exceeds a specific concentration. The run

length when sulfate is at or below 20 mg/L is 1500 bed volumes (BV). The run length when

sulfate is between 20 and 50 mg/L sulfate is 700 BV.

 2. The regeneration dose is also a critical element for the O&M costs. The salt dose for

regeneration was 10.2 lb/ft3.

 3. The second major component in the O&M costs is incremental labor cost. Incremental labor

is the labor associated with the additional maintenance that comes with a new process.

Adjustment in staffing or shifting of activities are not included in the incremental labor. The

incremental labor for the anion exchange is one hour per week plus three hours per

regeneration.

 4. The labor rate for small systems was $28/hour. The loaded labor rate for large systems was

$52/hour.

 5. O&M Costs were calculated using the data on regeneration dose, regeneration frequency,

labor hours per week and adjusted labor rates using the equation in the Very Small Systems

Document.

The capital and O&M cost curves for anion exchange are shown in figures 3-16 through 3-19. 
3-51

0

500

1000

1500

2000

2500

3000

3500

0 50 100 150 200 250 300 350

Sulfate Concentration (mg/l)

Bed Volumes to Arsenic Breakthrough

Upper curve is based on BV to 50% breakthrough

of arsenic. Lower curve is based on BV to 10% 

breakthrough (90% removal) of arsenic.

Both curves are based on trace levels of arsenic 

(< 1 mg/l) contamination in world average water

containing 20 mg/l chloride, and 200 mg/l

bicarbonate. Sulfate concentration is variable.

Figure 3-15

Bed Volumes to Arsenic Breakthrough as a Function 

of Sulfate Concentration (Clifford, 1995)
3-52

y = 344581x + 19546

y = -6278.7x2 + 319169x + 36758

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.01 0.1 1 10 100

Design Flow (mgd)

Capital Costs ($)

Figure 3-16

Anion Exchange (< 20 mg/L SO4)

Capital Costs
3-53

y = -67837x2 + 77544x + 4997.4

y = -1695.5x2 + 37350x + 16532

$100

$1,000

$10,000

$100,000

$1,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-17

Anion Exchange (< 20 mg/L SO4)

O&M Costs
3-54

y = 454220x + 19546

y = -9756.1x2 + 373847x + 73320

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.01 0.1 1 10 100

Design Flow (mgd)

Capital Costs ($)

Figure 3-18

Anion Exchange (20-50 mg/L SO4)

Capital Costs
3-55

y = -82812x2 + 94848x + 8449.9

y = -1843.1x2 + 48269x + 24614

$1,000

$10,000

$100,000

$1,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 3-19

Anion Exchange (20-50 mg/L SO4)

O&M Costs
3-56

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-mm 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. 

3.9.2 Ultrafiltration

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-mm prefiltration. UF has the benefit of being lower in both capital and O&M costs than highpressure membrane processes. Because UF is not an effective stand-alone technology for removal of

arsenic, capital and O&M cost estimates for UF are not provided in this chapter. 

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. Due to decreased removal efficiency when operated at higher recoveries, NF is not yet

demonstrated to be a reliable treatment for arsenic. Because NF is not an effective stand-alone

technology for removal of arsenic, capital and O&M cost estimates for NF are not provided in this

chapter. 

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
3-57

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-scalant before the membrane process. For the purpose of this analysis, costs were

not provided for a substantial pre-treatment system, other than the anti-scalant 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. Costs are not provided for RO because other options

are more cost effective and have much smaller waste streams. RO may be cost effective if removal

of other contaminants is needed and water quantity is not a concern. 

Both the VSS Model and the W/W Cost Model included cost estimation for RO. Since the

W/W 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 can be reduced by 50 percent. The membrane replacement portion of the O&M costs can also

be reduced by 50 percent to account for reductions in membrane costs. The W/W Cost Model for RO

was only valid up to a capacity of 200 mgd. The model also makes an assumption that recovery is 80

percent for systems of 1 to 10 mgd, and 85 percent for systems larger than 10 mgd. 

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 OHand 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.

Greensand filtration costs were not included in either the Water Model or the W/W Model. This

technology is considered to be a small systems technology and as a result costs were not estimated for

larger systems. This technology could be effectively operated in larger system sizes, but cost data are

not readily available. The key parameter is the ratio of source water iron and arsenic. If high
3-58

removals are needed the ratio needs to be at least 20:1 iron to arsenic. Costs are based on 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.

It is further assumed that corrosion control measures are not required because pH is not

affected by the process.

Figures 3-20 and 3-21 present cost estimates for removal of arsenic by greensand filtration.
3-59

y = 587584x0.838

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

Figure 3-20

Greensand Filtration

Capital Costs
3-60

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

y=0.0009x2 + 58921x + 7392.9

Figure 3-21

Greensand Filtration

O&M Costs
3-61

3.11 COMPARISON OF COSTS

The April 1999 and November 1999 drafts of the Technology and Cost Document compared

capital and O&M cost estimates in this chapter were compared with actual cost data and cost data

from other sources, including the 1993 draft Technology and Cost Document (Malcolm Pirnie, 1993).

There are numerous problems with the approach. The first is that there is not a consistent set of design

assumptions for a given technology. For example, without knowing the EBCT used in a full-scale

system, a comparison of capital costs is meaningless because the cost difference could be due to a

larger media volume. The cost estimates in Evaluation of Central Treatment Options as Small

System Treatment Technologies (SAIC, 1999) are not specific to arsenic removal. Differences could

be due to different assumptions about run length or process design. The same problem would apply

in making comparisons between the 1993 Technology and Cost Document and this final draft of the

Technology and Cost Document. Activated alumina and anion exchange do not rely on the cost models

used in that document because the models are inappropriate for arsenic removal costs. Therefore, this

document will not present any comparison of costs. 
3-62

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4-1

4.0 RESIDUALS HANDLING AND DISPOSAL ALTERNATIVES

4.1 INTRODUCTION

The implementation of any of the treatment technologies presented in Chapter 3 will result in

the generation of solid and/or liquid waste streams containing elevated levels of arsenic. It is the

purpose of this chapter to present the characteristics of the residuals generated by each of the treatment

technologies and to discuss appropriate handling and disposal options. 

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 those 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 quantity of residuals produced, which is dependent on

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 by a treatment process plays a significant role in determining

the handling and disposal method that will be utilized. For example, many handling methods which

are suitable for smaller systems are impractical for larger systems because of significant land

requirements associated with particular methods. As a result, larger systems that process residuals

on-site (as opposed to direct or indirect discharge) typically use mechanical methods to avoid large

outlays for land purchase, preparation, and maintenance. 

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

for large water systems. Transportation also can 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

greater maintenance than less complex methods.
4-2

4.l.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 first step in residuals processing, preceding

mechanical dewatering processes, evaporation ponds, and/or storage lagoons.

Filter backwash streams are high volume, low solids slurries generated during the cleaning

of granular filter media. Backwash volume depends on the number of filters used by a system and the

frequency with which they are cleaned. Typical backwash 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 characterized

by an average solids concentration of 0.5 to 2.0 percent (DPRA, 1993a).

When possible, backwash waters are recycled within the treatment process. In gravity

thickening, backwash waters are fed to a tank where settling occurs naturally. The sludges resulting

from gravity thickening are discharged and further treated for ultimate disposal, and the decant is

either recycled or discharged to a surface water or to a publicly-owned treatment works (POTW).

Gravity thickening reduces the quantity of water sent to waste due to backwashing, as well as the total

quantity of sludge generated (DPRA, 1993a). When recycling is not feasible, backwash waters may

be discharged to a surface water or a POTW, or may be treated by other mechanical or non-
4-3

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 and high O&M costs, compared to non-mechanical dewatering processes that can handle

similar waste volumes, e.g., storage lagoons. Due to their high costs, mechanical dewatering

processes are generally not suitable for application at very small water systems.

Filter presses have been used in industrial processes for decades, and their use has been

increasing in the water treatment industry over the past several years. These devices have been

successfully applied for the treatment of 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 may increase 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,

but have high capital costs and are labor intensive (DPRA, 1993a). Capital and O&M costs are

generally higher than those of comparable non-mechanical dewatering alternatives. As a result,

pressure filtration is most applicable for 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 on an influent solids concentration of 1 to 10 percent. Centrifugation is a

continuous process that requires very little time (8 to 12 minutes) to achieve optimal sludge solids

concentration. Centrifuges have low land requirements and high capital costs. They are more labor

intensive than non-mechanical alternatives, but less labor-intensive than filter presses. Again, due to

the associated capital and O&M requirements, centrifuges are more suitable for larger water systems.

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 (RO) and ion exchange (IX) (DPRA, 1993a). Ponds and drying beds are not

generally suitable for dewatering alum and lime sludges.
4-4

The use of evaporation ponds and drying beds involve the discharge of brine waste to a pond

for storage and dewatering via the natural process of evaporation. Typically, such ponds are designed

to have large surface areas to maximize evaporation of residual water by the sun and wind. Pond size

is determined by waste flow and storage capacity requirements.

As previously mentioned, evaporation ponds and drying beds are used primarily for brine

wastes generated by RO and IX. Such processes produce large volumes of liquid wastes

characterized by high levels of total dissolved solids (TDS), for which mechanical dewatering

processes, such as filter presses, are impractical. Depending on the solids concentration of the brine

waste stream, intermittent removal of solids may be required. For brines with a TDS content ranging

from 15,000 to 35,000 mg/L, solids will accumulate in an evaporation pond at a rate of 1⁄2 to 11⁄2

inches per year (DPRA, 1993a). When the depth of the solids in the pond reaches a predetermined

level, discharge to the pond is halted and evaporation is allowed to continue until the solids

concentration of the waste is suitable for disposal.

The use of evaporation ponds is extremely land intensive, requiring shallow basins with large

surface areas. This can be an important consideration in densely populated regions. Reverse osmosis

processes produce very large volumes of liquid wastes which increases land requirements and

ultimately construction costs. As a result, evaporation ponds may not be suitable for large water

systems utilizing RO. Evaporation ponds and drying beds have few O&M 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 be discharged to the

same pond for several years (assuming adequate total capacity) 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, lagoons, like evaporation ponds, are land intensive (DPRA, 1993a).

Storage lagoons are lined ponds designed to collect, retain, and dewater sludge for a predetermined

period of time. Dewatering occurs through the evaporation and decanting of the supernatant. Lagoon

size is determined based on the volume of sludge that must be handled and the desired storage time.

As with evaporation ponds, when a lagoon reaches its design capacity, solids can be removed by

means of heavy equipment and shipped for disposal.
4-5

Storage lagoons are best suited for dewatering lime softening (LS) process sludges, though

they have been applied with some success to coagulation/filtration (C/F) process sludges. Storage

lagoons can operate under a variety of sludge flows and solids concentrations, and chemical

conditioning of alum sludges is not required (DPRA, 1993a). Typically, lime sludges are discharged

to a lagoon with 3 percent solids, and are dewatered to 50 to 60 percent solids, whereas alum sludges

are typically characterized by 1 percent solids when first discharged to a storage lagoon and can be

dewatered to 7 to 15 percent solids (DPRA, 1993a). As evident from this data, alum sludges do not

typically dewater well in storage lagoons. When the top layer of an alum 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 for acceptance at most

landfills (DPRA, 1993a). Further, thickened alum sludges can be difficult to remove from lagoons,

often requiring 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 assist in the dehydration of alum sludges. 

4.3 DISPOSAL ALTERNATIVES

4.3.1 Direct Discharge

Direct discharge to a surface water is a method of disposal for some water treatment

byproducts. No pretreatment or concentration of the byproduct stream is necessary prior to discharge,

and the receiving water will naturally dilute the waste concentration and gradually incorporate the

sludge or brine (DPRA, 1993a).

However, the discharge of liquid residuals to surface waters is subject to compliance with and

permitting under the National Pollution Discharge Elimination System (NPDES). NPDES establishes

limits for specific contaminants (including total solids and arsenic) based on a variety of factors,

including ambient contaminant levels, low flow condition of the receiving water, and anticipated

volume of the proposed discharge. Most NPDES limits for solids discharge are around 30 mg/L

(AWWARF, 1998).
1

 Note that State and local limits may be more stringent than federal limits for specific contaminants,

including arsenic.

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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 Fg/L. If a

water source is also used as a source of drinking water, the arsenic limit is reduced to 0.0175 Fg/L;

although some States use the drinking water MCL for this purpose. These criteria are used by state

regulatory agencies as a basis for the determination of discharge limits for arsenic depending on the

classification of the receiving water.1 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 purchasing and installing the

necessary 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 successfully to dispose of alum and lime sludges, as well as brine streams generated by

systems using RO and IX (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 disposal option may require the installation of a conveyance system (i.e., piping) 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 LS sludges have also been successfully disposed of in this manner.

However, the POTW receiving such wastes 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.

The residuals generated from an arsenic treatment process may be classified as an industrial

waste since they contain contaminants, namely arsenic, which may impact the quality of the sludges

generated by the POTW. As a result, discharge to a POTW is only acceptable when arsenic

concentrations fall within the Technically Based Local Limits (TBLL) established by the POTW under
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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 municipal wastewater. Background levels are usually

measured in domestic wastewater. The existing TBLL would be based on the untreated drinking water

quality (higher arsenic). Removing arsenic from drinking water would change the background level

in the municipal wastewater. The TBLL would likely be revised since the background level has been

lowered. The revised TBLL would be used to determine if the brine stream from the drinking water

process could be discharged to the POTW. One approach to evaluating if indirect discharge is an

option is to perform a mass balance on arsenic based on the treated and untreated conditions. The

critical factor in determining the arsenic increase at the POTW will be the volume of water lost

between the drinking water treatment plant and the POTW. If there were no water lost between the

drinking water treatment plant and the POTW, then there would be no arsenic increase to the POTW.

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 of the residual exceeds the Clean Sludge Limit of 41 mg/kg, the biosolids may

be land applied, but the quantity will be limited to a total cumulative arsenic loading of 41 kg per

hectare (36.6 lb/acre). As a result, most TBLLs are based on the Clean Sludge criterion (AWWARF,

1998). This source appears to overstate the impact of the arsenic concentration in sludge and the

ability to land apply the biosolids. Biosolids with concentrations between 41 and 75 mg As/kg

biosolids can be land applied as long as the arsenic accumulation is tracked. The lifetime

accumulation is 41 kg As/hectare of land. Using the maximum arsenic concentration of 75 mg/kg

biosolids, over 500,000 kg of biosolids would need to be applied to one hectare to exceed the limit,

which is unlikely.

To illustrate the impact of Part 503 regulations on the development of an arsenic TBLL,

consider a limit of 41 mg/kg for the land application of biosolids (AWWARF, 1998). 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
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equates to a total (municipal and industrial) influent concentration of around 13 Fg/L (AWWARF,

1998). As a result, if a water system is located in an area that has a background arsenic concentration

near 13 Fg/L, it may not be permitted to discharge to the local POTW. There are several problems

with the above example. The first is that the background concentration would be reduced because

arsenic is being removed from drinking water. Therefore, the TBLL would need to be revised. The

second is that 41 mg/kg is being used as the upper limit for land application when the actual upper

limit is based on 75 mg As/kg biosolids since arsenic accumulation is unlikely to exceed the lifetime

limit. The key variables to determine if arsenic in the POTW biosolids would exceed the land

application limit are: existing background arsenic concentration, water loss between drinking water

treatment plant and POTW, arsenic removal efficiency of the POTW, and the volume of biosolids

generated by the POTW. Land application restrictions on the biosolids would only likely be an issue

when background arsenic levels are high, water loss is 50% or higher between the drinking water

treatment plant and the POTW, the POTW has a high arsenic removal efficiency and only generates

a low volume of biosolids. Thus, arsenic accumulation in biosolids is unlikely to be as significant

a restriction as total dissolved solids (TDS) increases due to the salt used for regeneration for anion

exchange.

Anion exchange is regenerated using sodium chloride. Typical regenerations will use 10.2

lb/cubic feet of resin. Even though the brine stream is small in volume, the TDS concentration of the

brine stream will be very high. As a result, the POTW discharge may increase the TDS content of the

receiving waters. Areas of the country with high TDS naturally or limited quantities of water may find

any increase in TDS to be unacceptable. In areas where small increases in TDS can be

accommodated, the water loss between the drinking water treatment plant and the POTW and the size

of the brine stream compared to the total volume of water treated at the POTW will affect the

magnitude of the TDS increase.

As with direct discharge, 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. Other 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 placed on the local POTW.
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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 on several variables, including soil and sludge chemistry and 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 up to 41 mg/kg can be land applied without tracking arsenic accumulation. When the

concentration is in the range of 41 to 75 mg/kg, sludges can be land applied if arsenic accumulation

is tracked and does not exceed 41 kg As/hectare. Due to the possibility of arsenic absorption by

vegetation, application of sludges to non-food chain fields is preferred. Land application is also

limited by the availability of land. In areas where grassland, farmland, or forested land is

unavailable, transportation costs can significantly affect the overall cost effectiveness of this disposal

option.

Land application can serve as a means of final disposal of LS, and to a lesser degree C/F,

sludges. Lime sludges can have a beneficial impact on farmland by neutralizing soil pH, replacing

the use of 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

governing the types of wastes that 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). Federal requirements include restrictions on

location, operation and design criteria, groundwater monitoring requirements, corrective action

requirements, closure and post-closure requirements, and financial assurance.
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Landfill disposal requires that residuals be in a solid form and contain no free liquids

(typically determined by the Paint Filter Liquids Test, SW-846, Method 9095). 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 meeting even one of these criteria is considered hazardous. For drinking water

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 toxicity characteristic (TC) regulatory level

for arsenic is 5 mg/L, which is 100 times the interim primary drinking water maximum contaminant

limit (MCL) of 50 Fg/L. The TC regulatory levels for other inorganic contaminants were also based

on 100 times their interim primary drinking water MCL. The MCLs for many of these inorganic

contaminants have been revised since 1990 whereas the TC levels have not changed. Thus, lowering

the arsenic MCL does not mean that the TC regulatory level will be lowered. A separate rulemaking

would be required for the reduction of the TC level. 

Many water treatment facilities currently dispose of their waste residuals in commercial or

public 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 associated with the development of a sanitary landfill (DPRA,

1993a). Monofills control the type of waste that they accept more strictly and also limit potential

future liabilities. However, it is unlikely that monofills will be utilized for arsenic treatment

residuals.

4.3.5 Hazardous Waste Landfill Disposal

Water treatment residuals containing arsenic which fail the TCLP test for toxicity must be

disposed of 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 the disposal of arsenic-containing residuals in a hazardous

waste landfill is the presence of free liquids. If any water treatment sludge contains free liquids,
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usually 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.

Hazardous waste 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 a treatment residual is 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, the

generation 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. Due to the extremely high costs

associated with this disposal method, it was assumed for the purpose of estimating costs for the

Arsenic Rule that water systems would not select treatment processes that would require disposal of

waste residuals as hazardous waste to meet the revised MCL.

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, typically

characterized by 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 in

previous sections of 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 (mgd).

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 previous sections of this chapter. As with sedimentation sludges, discharge of filter
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backwash streams to a sanitary sewer is generally only an option for treatment plants with an average

flow of less than 10 mgd.

Selection of Handling and Disposal Options

As previously discussed, C/F blowdown and filter backwash are high volume waste streams

characterized by low solids content. Gravity thickening may be used to increase the solids content of

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 a land-intensive

disposal option, they may not be applicable for large water systems.

Disposal of C/F residuals containing arsenic is largely dependent on influent arsenic

concentration, coagulant dose and suspended solids content. Disposal by direct discharge to a surface

water is not typically appropriate. 

Depending on the arsenic concentration of C/F sludges, land application may be a suitable

disposal method. As discussed in section 4.3.2, total arsenic cannot exceed 41 mg/kg if sludges are

to be applied with no restrictions. Sludges with arsenic concentrations between 41 and 75 mg/kg may

be land applied provided that the total loading does 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 of in a sanitary landfill. Otherwise, residuals must be disposed of in a

hazardous waste landfill. A 1992 study (Bartley et al., 1992) found that the waste sludges from C/F

plants would pass the TCLP test. In this study, samples were taken from the waste sludges generated

by four different water treatment plants and subjected to the TCLP test. One of these systems relied

on C/F, and another used both C/F and lime softening processes. These two systems treated raw

waters characterized by arsenic concentrations averaging 1.1 mg/L and less than 0.001 mg/L,

respectively. The results of the TCLP tests conducted on the waste residuals from both systems were

below 0.040 mg/L (range: 0.007 to 0.039 mg/L)–well below the current criterion for treatment as

hazardous waste (5.0 mg/L).

A recent study (EPA, June 2000), examined the characteristics of the waste sludges generated

by two C/F plants (Plant A and Plant B). The sludge from Plant A was generated from backwashing

anthracite coal/pea gravel filters, while that from Plant B was generated as a result of sedimentation

in primary and secondary clarifiers and from filter backwashing. Both sludges passed the TCLP test
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for arsenic toxicity by a substantial margin (ranges: less than 0.050 to 0.106 mg/L and 0.058 to 0.160

mg/L, respectively) and would not have exceeded the strict TCLP limits established by California for

nonhazardous waste. However, the sludge from Plant A would violate the soluble threshold limit

concentrations (SLTCs) established by California for arsenic and copper. As a result, the sludge

generated by this treatment plant would be considered a hazardous waste, requiring appropriate

handling and disposal.

Tests conducted by the University of Colorado confirm the results found in these two studies,

indicating 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 or exceeds another

regulatory limit established for water treatment sludges by individual States.

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 only about 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 mgd.

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 section 4.2.2. As for 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

mgd.

Selection of Handling and Disposal Options

As previously discussed, enhanced coagulation blowdown is characterized by high volume

and low solids content. Typical solids contents range from 0.5 to 2.0 percent, depending on coagulant

type. Gravity thickening may be used to increase the solids content of enhanced C/F sludges prior to
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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 may not be applicable for large water systems since they are

land intensive.

Disposal of enhanced coagulation arsenic residuals is largely dependent on influent arsenic

concentration, coagulant dose, and suspended solids content. Disposal by direct discharge to a surface

water is not typically appropriate. 

Depending on the arsenic concentration of C/F sludges, land application may be a suitable

method of disposal. Total arsenic should not exceed 41 mg/kg if sludges are to be land-applied

without restrictions. Sludges with arsenic concentrations between 41 and 75 mg/kg may be land

applied provided that the total loading does 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 of in a sanitary landfill. Otherwise, residuals must be

disposed of in a hazardous waste landfill. Note that hazardous waste disposal is unlikely since tests

conducted by the University of Colorado indicate that enhanced coagulation sludges will pass the

TCLP test (AWWARF, 1998). Due to its high cost disposal to a hazardous waste landfill should only

be relied on 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 used in this

process. Sludge production is also affected by the suspended solids content of the raw water.

Systems are unlikely to install direct filtration solely for arsenic removal. If this technology were

installed, the disposal alternative discussed for coagulation/filtration would be appropriate for this

technology.

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(CH)3 precipitate.
4-15

The solids content of the backwash from coagulation assisted microfiltration processes were found

to be less than 0.5 percent in one study (Clifford et al, 1997). 

Selection of Handling and Disposal Options

The wastes from coagulation assisted microfiltration processes will consist of a very dilute

slurry. Gravity thickening may be used as to increase the solids content of the sludge prior to the use

of 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 means of handling the residuals generated by coagulation assisted microfiltration. Both

require little oversight and maintenance, but are land intensive. As such, they may not be appropriate

for large systems. A thorough comparison of handling options should be conducted to select the most

cost effective method.

Land application may be a suitable disposal method for sludges from coagulation assisted

microfiltration processes. As discussed in section 4.3.2, total arsenic cannot exceed 41 mg/kg if

sludges are to be applied with no restrictions. Sludges with arsenic concentrations between 41 and

75 mg/kg may be land applied provided that the total loading does not exceed 41 kg 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. However, tests conducted by Clifford (Clifford,

1997) and the University of Colorado (AWWARF, 1998) indicate that the sludges from this treatment

process will pass the TCLP test for arsenic toxicity by a considerable margin, making it unlikely that

hazardous waste disposal will be necessary. 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 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 the use of LS for the removal of non-carbonate hardness.

LS plants typically produce 1,000 to 8,000 pounds of solids per million gallons of water

treated depending on the hardness of the influent water (AWWARF, 1998). Arsenic concentrations
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of these sludges, however, are generally lower than those found in C/F sludges due to the increased

volume of solids produced.

Selection of Handling and Disposal Options

LS blowdown is slightly more dense than C/F blowdown. Typical solids content ranges from

1.0 to 4.0 percent, depending on raw water hardness. Gravity thickening may be used to pretreat LS

sludges prior to their 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.

Land application of LS treatment sludges is one possible disposal alternative. As discussed

in section 4.3.2, total arsenic cannot exceed 41 mg/kg if sludges are to be applied with no restrictions.

Sludges with arsenic concentrations between 41 and 75 mg/kg may be land applied provided that the

total loading does not exceed 41 kg per hectare. 

LS sludges will require dewatering prior to landfill disposal. If the residuals pass the TCLP

test they may be disposed of in a sanitary landfill. Otherwise, residuals must be disposed in a

hazardous waste landfill. However, hazardous waste landfill disposal should only be used as a last

resort if waste fails the TCLP test which is unlikely given the findings of available studies. Tests

conducted by the University of Colorado indicate that LS sludges will pass the TCLP test (AWWARF,

1998). This finding is supported by the findings of Bartley, et al. (1992).

In the Bartley study, samples were taken from the waste sludges generated by four different

water treatment plants and subjected to the TCLP test. Two of these systems relied on LS, and another

used both C/F and LS processes. All three of these systems treated raw waters characterized by

arsenic concentrations averaging less than 0.001 mg/L. The results of the TCLP tests conducted on

the waste residuals from these systems ranged from 0.007 to 0.039 mg/L–significantly below the

current criterion for treatment as hazardous waste (5.0 mg/L).

4.4.6 Enhanced Lime Softening

Enhanced LS is a modified LS process wherein lime dosage is increased. Residuals generated

as a byproduct of this process are similar to those generated by a typical LS treatment process. The

quantity of sludge produced is a function of water hardness. As with LS, the use of enhanced LS for
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carbonate hardness removal produces approximately twice the amount of solids per pound of hardness

removed as the use of this process for non-carbonate hardness removal.

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.

Selection of Handling and Disposal Options

Typical solids content of enhanced LS blowdown range from 1.0 to 4.0 percent, depending on

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.

Land application of treatment sludges is a possible disposal alternative. As discussed in

section 4.3.2, total arsenic cannot exceed 41 mg/kg if sludges are to be applied with no restrictions.

Sludges with arsenic concentrations between 41 and 75 mg/kg may be land applied provided that the

total loading does not exceed 41 kg per hectare.

Sludges will require dewatering prior to landfill disposal. If the residuals pass the TCLP test

they may be disposed of in a sanitary landfill. Otherwise, residuals must be disposed of in a

hazardous waste landfill. However, as with LS sludges, such treatment is not likely to be necessary,

and should only be used as a last result.

Tests conducted by the University of Colorado indicate that enhanced LS sludges will pass the

TCLP test (AWWARF, 1998). Additional tests conducted by the University of Colorado confirmed

that the treatment sludge from enhanced LS treatment will typically pass the TCLP (results: 0.0009

to 0.0284 mg/L; Amy, 1999). A case study presented by Sorg (2000) lends further strength to the

expectation that the residuals from this treatment process may be treated as nonhazardous.

Sludge from the enhanced LS plant in a recent EPA study was collected, sent to a slurry tank,

and then to a sludge lagoon (EPA, June 2000) . Approximately once every 2 years, or as required,

the sludge lagoon was dewatered using an underdrain system. Arsenic was not detected in the

leachate at any of the sampling locations used in the study (detection limit: 0.05 mg/L). As a result,
4-18

following dewatering, the dry sludge from this plant could be removed and applied to local farm fields

as a beneficial amendment.

4.4.7 Ion Exchange

Ion exchange (IX) involves the use of a synthetic resin in the 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 chloride per equivalent of resin, i.e., 10.2 pounds of salt per cubic foot of resin.

Regeneration requires approximately 2.8 bed volumes (BV) of brine and 1.2 BV displacement rinse.

Therefore, 4 to 5 BV of waste are produced per regeneration cycle. (AWWARF, 1998). The rinse

volume in another study was 5 BV (Clifford et al, 1997). Using this as an upper bound, the waste

volume could be as high as 7.8 BV.

The other important factor is the arsenic concentration in the waste brine. An EPA study (EPA

October 2000) which found that arsenic concentrations during brine regeneration ranged from 1.83

to 38.5 mg/L (average: 16.5 mg/L). The wastewater generated during the other three steps of the

regeneration process (i.e., backwash, slow rinse, and fast rinse) evaluated by this study were

characterized by lower arsenic concentrations (0.0594, 1.332, and 0.108 mg/L, respectively).

Therefore, if these waste streams were combined prior to disposal, the overall arsenic concentration

of the brine would be below that observed during the brine regeneration step.

Selection of Handling and Disposal Options

The November 1999 Technology and Cost Document listed three mechanisms to dispose of

the brine stream used for regeneration. The options were: sanitary sewer, evaporation pond, and

chemical precipitation. Many comments on the proposed rule were considered with waste streams

being considered hazardous waste. Waste streams containing less than 0.5% solids are evaluated

against the toxicity characteristic directly to determine if the waste is hazardous. Arsenic in the

regeneration brine will likely exceed 5 mg/L for most systems with arsenic above 10 Fg/L and sulfate

below 50 mg/L. Since the brine stream would likely be considered hazardous, the evaporation pond

and the chemical precipitation options were eliminated as options for disposal of anion exchange

wastes. Discharge to a sanitary sewer was retained because domestic sewage and any mixture of

domestic sewage and other wastes that pass through a sewer system to a publicly-owned treatment

works (POTW) for treatment is excluded from being considered solid waste (40 CFR 261.4).
4-19

Domestic sewage means untreated sanitary wastes that pass through a sewage system. Discharges

meeting the above criteria are excluded from regulation as hazardous waste.

Although the brine stream may be deemed a hazardous waste, such characterization would not,

in itself, prevent disposal to a POTW since wastes that passes through a sewer system to a POTW are

exempt from RCRA regulation. Discharge to the sanitary sewer may be limited by TBLLs for arsenic

or TDS. The TDS is the more likely limiting factor since the arsenic TBLL would need to be

recalculated since the background arsenic will change. The only net increase in arsenic is related to

the water loss from the drinking water treatment plant to the POTW. TDS is different. TDS is being

added to the waste stream. Due to the potential for an increase in total dissolved solids, anion

exchange would be favored in areas other than the southwest where the volume of brine is very small

relative to the total volume of wastewater being treated at the POTW. 

Although one small system was reported to have disposed of its waste brine to the facility’s

septic tank (EPA October 2000), such disposal is likely to be regulated by local and State regulations

and therefore was not considered in the cost analysis for the final rule. 

4.4.8 Activated Alumina

Activated alumina (AA) can be operated with or without pH optimization and regeneration.

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, 1998).

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 (approximately 42 BV per regeneration cycle). Arsenic

is strongly adsorbed to the media, so only about 50 - 70% of the adsorbed arsenic is removed even

using a strong base.. The brine stream produced by the regeneration process then requires disposal.

The November 1999 Technology and Cost Document listed discharge to a sanitary sewer as the

disposal mechanism for the brines. It was noted that TBLLs for arsenic might restrict discharge of

brine streams to the sanitary sewer. Since activated alumina run lengths are much longer than anion

exchange, the arsenic concentrations in the brine stream would likely be much higher even with

incomplete regeneration. Regeneration of activated alumina media is not recommended even for

larger systems because disposal of the brine may be difficult, the regeneration process is incomplete

which reduces subsequent run lengths, and for most systems it will be cheaper to replace the media

rather than regenerate it. As a result, regeneration of AA was not considered for the final rule.
4-20

Selection of Handling and Disposal Options

Residuals from systems operating on a ‘throw away’ basis, will not face the difficulties of

those systems that regenerate since they need only dispose of spent media, rather than both

concentrated brines and spent media. These systems can dispose of process residuals in a nonhazardous waste landfill. Spent alumina will have to pass the TCLP test before disposal at a sanitary

landfill is possible. In a recent study (Wang et al., 2000), TCLP tests were performed on samples of

activated alumina taken from two different water systems. Both of these systems had similar set-ups:

four tanks of activated alumina with two parallel sets of two tanks in series (a lead or ‘roughing’ tank

followed by a ‘polishing’ tank). Each tank contained approximately 10 cubic feet of activated

alumina. During operation, the flow rate through each parallel series was 5 gallons per minute. This

resulted in a hydraulic loading rate of 1.59 gallons per minute per square foot and an empty bed

contact time (EBCT) of 15 minutes. Over the course of the study, the influent concentration of arsenic

ranged from 0.053 to 0.087 mg/L (average: 0.062 mg/L) and 0.021 to 0.076 mg/L (average: 0.049

mg/L) for the two systems. Both systems consistently removed greater than 92 percent of influent

arsenic throughout the course of the study. The media from each of the roughing tanks was replaced

about every 18 months. At the time of media replacement, samples of spent media from different parts

of the roughing tanks (i.e., the top, middle, and bottom) were taken and subjected to the TCLP test.

In all cases, the samples had TCLP test results of less than 0.05 mg/L, well below the statutory limit

of 5.0 mg/L. Based on these findings and the fact that activated alumina binds most tightly to arsenic

at the pH at which the TCLP is conducted, it is unlikely that the spent media will be classified as

hazardous. 

4.4.9 Nanofiltration

NF membranes are primarily used for softening, removing the larger, divalent ions associated

with hardness (AWWARF, 1998). The recovery rate is dependent upon the source water quality.

Typical recovery rates of this technology approach 85 percent. The reject stream from NF contains

elevated levels of arsenic and other contaminants that are removed from the source water by the NF

membranes. As discussed in Chapter 3, nanofiltration is not yet a proven reliable technology that can

be operated at realistic recoveries. If data are generated to show it can be reliably operated at high

recoveries, then the disposal options would be similar to those for reverse osmosis.
4-21

4.4.10 Reverse Osmosis

Reverse osmosis (RO) membranes will remove much smaller ions typically associated with

TDS (AWWARF, 1998). The recovery rate for RO membranes is dependent on the source water

quality but typically falls between 30 and 85 percent. The reject stream from this process contains

elevated levels of arsenic and other contaminants. While reverse osmosis is an effective removal

technology, it is unlikely to be selected solely for arsenic removal because other options are more cost

effective for arsenic removal and do not reject a large volume of water. RO may be cost effective

options if removal of other contaminants is needed and water quantity is not a concern.

The disposal technologies were presented in the November 1999 Technology and Cost

Document. They were: direct discharge, indirect discharge, and chemical precipitation/non-hazardous

landfill disposal. Since the recovery rate is typically less than 85% percent, reverse osmosis reject

streams can be characterized as high volume with lower arsenic concentrations compared to the brine

streams from anion exchange and activated alumina. Arsenic concentrations should not exceed the TC

regulatory level since the maximum concentration effect is less than 6 to 1.

Selection of Handling and Disposal Options

Direct discharge of RO residuals to a receiving surface water is one possible disposal

alternative. The large volume of the waste stream makes this alternative unlikely except for small

systems (which produce a small waste volume) or systems located near the coast.

Indirect discharge to the POTW is also an option. Similar to anion exchange, the existing

TBLLs would likely need to be modified because the background arsenic will change. The water loss

between the drinking water plant and the POTW will be the critical factor in determining the arsenic

increase to the POTW. The volume of the reject stream compared to the total volume at the POTW

will also be a critical factor in determining if the POTW would accept this waste stream. Systems

should check with the POTW before selecting this option.

Chemical precipitation of reverse osmosis streams is likely to be cost prohibitive. Reverse

osmosis is already a fairly expensive treatment technology compared to other options and chemical

precipitation of the large volume waste stream would significantly increase those costs. If direct

discharge and indirect discharge are unavailable, systems should consider another treatment

technology.

Costs are not presented for reverse osmosis waste disposal because it was not used to develop

national costs and other options are more cost effective and produce much smaller waste streams.
4-22

4.5 RESIDUALS HANDLING AND DISPOSAL COSTS

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 four water treatment processes: (1) coagulation assisted

microfiltration, (2) ion exchange, (3) activated alumina, and (4) greensand filtration. The disposal

options assumed for each of these treatment processes are shown below:

# Ion Exchange:

- POTW (indirect) discharge

# Coagulation Assisted Microfiltration:

- Mechanical dewatering and nonhazardous landfill disposal

- Non-mechanical dewatering and nonhazardous landfill disposal

# Activated Alumina:

- Nonhazardous landfill disposal (systems operating without regeneration)

# 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, 1993a and 1993b). Capital and O&M cost estimates for

the selected handling and disposal options are presented in Figures 4-1 through 4-14. Capital cost

estimates for nonhazardous landfill disposal are assumed to be zero. As a result, capital cost

estimates are not provided when landfill is the sole disposal technology under consideration (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

nonhazardous landfill disposal. Land costs are excluded from all capital cost estimates.
4-23

y = 21.47x + 5198

$100

$1,000

$10,000

$100,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

y = 3955 y = 5085

Figure 4-1

Anion Exchange (< 20 mg/L SO4)

Waste Disposal Capital Costs
4-24

$100

$1,000

$10,000

$100,000

$1,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

y = 3429x + 375

Figure 4-2

Anion Exchange (< 20 mg/L SO4)

Waste Disposal O&M Costs
4-25

y = 0x + 3955

y = 5085

y = 21.47x + 5198

$100

$1,000

$10,000

$100,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

Figure 4-3

Anion Exchange (20-50 mg/L SO4)

Waste Disposal Capital Costs
4-26

y = 7348x 

$1

$10

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 4-4

Anion Exchange (20-50 mg/L SO4)

Waste Disposal O&M Costs
4-27

y = -692793x2 + 455329x + 26748

y = 211627x + 42043

y = -1644.1x2 + 150402x + 157575

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

Figure 4-5

Coagulation Assisted Microfiltration (Mechanical)

Waste Disposal Capital Costs
4-28

y = -3476860x2 + 684838x + 5839

y = 25165x + 36857

y = -359.66x2 + 46936x + 21793

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 4-6

Coagulation Assisted Microfiltration (Mechanical)

Waste Disposal O&M Costs
4-29

y = 3069360x - 790

y = 1749352x + 108017

y = 1627970x + 326504

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.01 0.1 1 10 100 1000

Design Flow (mgd)

Capital Costs ($)

Figure 4-7

Coagulation Assisted Microfiltration (non-mechanical)

Waste Disposal Capital Costs
4-30

y = 18812x2 + 4686x + 2124

y = 111819x - 6950

y = 16.966x2 + 60792x + 28760

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

$100,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 4-8

Coagulation Assisted Microfiltration (non-mechanical)

Waste Disposal O&M Costs
4-31

y = 7568x

$1

$10

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 4-9

Activated Alumina (pH 7 - 8)

Waste Disposal O&M Costs
4-32

y = 14555x

$1

$10

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 4-10

Activated Alumina (pH 8 - 8.3)

Waste Disposal O&M Costs
4-33

y = 3276x

$1

$10

$100

$1,000

$10,000

$100,000

$1,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 4-11

Activated Alumina (pH Adjusted to 6.0, 23,100 BV)

Waste Disposal O&M Costs
4-34

y = 4915x 

$1

$10

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.001 0.01 0.1 1 10 100 1000

Average Flow (mgd)

O&M Costs ($)

Figure 4-12

Activated Alumina (pH Adjusted to 6.0, 15,400 BV)

Waste Disposal O&M Costs
4-35

$100

$1,000

$10,000

$100,000

0.01 0.1 1

Design Flow (mgd)

Capital Costs ($)

y = 3955

Figure 4-13

Greensand Filtration

Waste Disposal Capital Costs
4-36

y = 7548x + 424

y = 7548x + 1130

$100

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

0.001 0.01 0.1 1

Average Flow (mgd)

O&M Costs ($)

Figure 4-14

Greensand Filtration

Waste Disposal O&M Costs
5-1

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 where 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. Wholehouse, 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 500 or more, but less than 3,300; and

# Population of 3,300 or more, 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.”

Research has shown that POE and POU devices can be an effective means of removing arsenic

from potable water (Fox and Sorg, 1987; Fox, 1989). Influent arsenic concentration and other source
5-2

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, H2AsO4

2- and HAsO4-

. 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.

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.

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
5-3

chromatographic 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.
5-4

TABLE 5-1

Source Water Summary - Point-of-Use Case Studies

Source Water Characteristic Concentration1

Alkalinity 56 - 206

Arsenic <0.005 - 1.16

Calcium 8.9 - 22

Chloride <10

Fluoride 0.6 - 5.2

Iron <0.1 - 2.5

Magnesium 5.3 - 10.6

Manganese <0.6

pH 7.4 - 8.3

Silica NA

Sodium 4.4 - 62

Sulfate <15

Total Dissolved Solids <1,500

Total Hardness 109 - 547

Turbidity 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

Arsenic 

Removal

(%)

Quantity of Water

Treated When MCL

Exceeded* (gal)

Time On-Line Until

MCL Exceeded*

(days)

Reverse Osmosis

(low pressure) 50-86 50-332 14-225

Reverse Osmosis

(high pressure) 50-80 684* 650**

Activated Alumina NA 0-18,557** 0-1,226**

Ion Exchange NA 0-20,935** 0-1,471**

* Based on EPA DWRD study conducted in Fairbanks, Alaska and Eugene, Oregon with arsenic concentrations ranging from 0.026 to 1.16 mg/L. (Fox,

1989)

** Taken off line before MCL exceeded.

NA = not applicable. Removal percentage does not adequately express unit performance for these technologies because run length is the determining factor

when attempting to achieve a desired removal percentage. 
5-5

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 on start-up, achieving 60 to 80

percent removal of arsenic. Over time, however, the arsenic removal efficiency decreased to 50

percent 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 percent removal,

but with high influent arsenic concentrations, much higher removal efficiency is necessary to achieve

the MCL. The high-pressure unit (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 IX 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 IX beds with the exception that they were filled with

granular activated alumina (Alcoa-F1) rather than resin. Over the course of this study, the AA beds

effectively reduced arsenic levels; however, the media required pre-treatment to reduce the pH of the

AA to 5.5 - 6.0 (the pH at which AA most effectively removes arsenic). Pre-treatment of the AA bed

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 longer than one year.

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; 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
5-6

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.

During the study, observed arsenic concentrations of 0.068 mg/L to 0.02 mg/L were

consistently reduced to less than the detectable limit (0.005 mg/L) by the POU RO units. Other

contaminants were also effectively removed, including manganese (80 percent), iron (85 percent), and

TDS (95 percent). Based on the manufacturer’s literature, it appeared that the units were operating

at an approximate recovery rate of 25 percent, 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 wellhead. As a result, a carbon prefilter 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 on 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 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.
5-7

Manufacturer and laboratory data suggest greater than 95 percent removal of arsenate by RO

systems, and slightly less (75 percent) removal of arsenite. Field studies indicate that greater than 50

percent removal is possible, but data are inconclusive much beyond those levels. 

5.4.1 Cost Estimates

The EPA document, Cost Evaluation of Small System Compliance Options - Point-of Use

and Point-of-Entry Treatment Units (Cadmus Group, 1998), was used to estimate POE and POU RO

treatment costs. Costs are presented in Figures 5-1 and 5-2 for total capital costs and annual O&M

costs for POU RO treatment, respectively. The cost curves are based on 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 percent discount for 10 or more

units, 15 percent 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.

POE RO is not considered a compliance technology because it could create corrosion control

problems. Furthermore, water recovery would not be any higher than observed for central RO

treatment, so concerns regarding water quantity that arise with central treatment would also apply to

POE treatment. Therefore, POE RO treatment option is not considered in the final rule.
5-8

y = 864.66x0.9261

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

10 100 1000 10000

Households

Capital Costs ($)

Figure 5-1

POU Reverse Osmosis

Total Capital Costs
5-9

y = 266.92x0.9439

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

10 100 1000 10000

Households

O&M Costs ($)

Figure 5-2

POU Reverse Osmosis

O&M Costs
5-10

5.5 ION EXCHANGE

Ion exchange (IX) has been used effectively for the removal of arsenic. There are two types

of IX 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 as the basis for estimates of POE and

POU treatment costs. The following assumptions were made to estimate the cost of POE and POU IX:

# 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 percent discount for 10 or more

units, vendor retains 30 percent profit on more than 100 units.
5-11

# 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.

Due to concerns about chromatographic peaking and run lengths that typically would be less

than 6 months, POU anion exchange is not listed as a compliance technology in the final rule (Kempic,

2000). POE IX also may present problems since the brine waste from the regeneration process is

characterized by high TDS. In addition, POE run lengths were determined to be only a few months,

so frequent regeneration would be required (Kempic, 2000). Therefore, POE anion exchange is also

not listed as a compliance technology in the final rule.

5.6 ACTIVATED ALUMINA

Activated alumina (AA) can be used to remove inorganic contaminants, including arsenic, from

source water. Contaminant removal occurs as contaminants are exchanged with the hydroxide ions

on the alumina surface. Since AA is an amphoteric material, its properties are closely linked to the

pH of the ambient environment. At high pH (i.e., pH above 8.2), AA will act as a cationic exchanger,

while at lower pHs, it will perform as an anionic exchanger. However, it is important to note that AA

is rarely used at high pH (>8.3) 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 with central treatment will need some type of pH adjustment to accommodate this

requirement. For POE and POU systems this can not be accomplished because of practical reasons.

5.6.1 Cost Estimates

The Cadmus Group (1998) document was again used as the basis to develop estimates for POE

and POU treatment costs. Costs are presented in Figures 5-3 for total capital costs and 5-4 for annual

O&M costs. The cost curves are based on 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).
5-12

# 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 percent discount for 10 or more

units, vendor retains 30 percent 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.

Arsenic removal by AA is very sensitive to pH, and finished water pH typically will be higher

than the optimal required for AA. A finished water pH for many systems would be in the range of pH

7 to pH 8. Using data on activated alumina run length and pH, it was determined that viable run

lengths for POE AA were likely only when the finished water pH was at or below pH 7.5 (Kempic,

2000). Even in this pH range, the media may need to be replaced more frequently than once a year,

which would make the option very expensive especially compared to the POU AA option. The run

length data used for this analysis were from a site with very little competing ions (Simms and Azizian,

1997). Studies at other sites with higher levels of competing ions have much lower run lengths

(Clifford et al, 1998). Based on the limited finished water pH range where POE AA might be

effective and that the POU media needs replacing much less frequently due to lower water demand,

POE AA has not been listed as a compliance technology. POE devices that utilize media that are less

sensitive to pH adjustment may be listed as compliance technologies in the future once data on their

performance is generated.

However, because the volume of water treated by POU AA is much smaller, POU units have

typically have longer run lengths and require less frequent replacement than POE units. Therefore,

the costs of POU AA treatment were considered in the final rule. The cost of replacement for POU

AA is also much lower than POE AA.
5-13

y = 296.89x0.9257

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

10 100 1000 10000

Households

Capital Costs ($)

Figure 5-3

POU Activated Alumina

Total Capital Costs
5-14

y = 412.61x0.9376

$1,000

$10,000

$100,000

$1,000,000

$10,000,000

10 100 1000 10000

Households

O&M Costs ($)

Figure 5-4

POU Activated Alumina

O&M Costs
6-1

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WMA, Inc. (1994). Assembly of Compliance Decision Tree for Arsenic, Correspondence with Ms.

Heather Shank-Givens, EPA, Office of Ground Water and Drinking Water, November 1994.
6-10

WMA, Inc. (1995). Arsenic: Impact Summaries (EPA Contract 68-C3-0368), Correspondence with

Ms. Heather Shank-Givens, EPA, Office of Ground Water and Drinking Water, January 1995.

Waypa, J.J., M. Elimelech and J.G. Hering (1997). “Arsenic Removal by RO and NF Membranes,”

J. AWWA, 89:10:102.

Westerhoff, G. and Z.K. Chowdhury (1996). “Water Treatment Systems,” Chapter in Water Resources

Handbook, L.M. Mays ed., McGraw Hill, New York, 1996.

Wiesner, M.R., S. Sethi, J. Hackney, J. Jacangelo and J.M. Laine (1994). “A Comparison of Cost

Estimates for Membrane Filtration and Conventional Treatment,” J. AWWA, 86:12:33-41.

Wilkie, J.A. and J.G. Hering (1998). “Rapid Oxidation of Geothermal Arsenic (III) in Streamwaters

of the Eastern Sierra Nevada,” Environmental Science and Technology, 32:5:657-662.
Appendix A Very Small Systems Capital Cost

Breakdown Summaries

Table A1 - VSS Document Capital Cost Breakdown for Membrane Processes

Component Capital Cost 

Factor

Percent of Total 

Capital Cost

Capital Cost 

Breakdown 

Category

Equipment 1.0000 56.97% p

Installation 0.2500 14.24% c

Sitework/Interface Piping 0.0750 4.27% c

Standby Power 0.0625 3.56% c

OH&P 0.1665 9.49% e

Legal & Admin 0.0416 2.37% e

Engineering 0.1596 9.09% e

Contigencies 0.0000 0.00% c

Total 1.7552 100.00%

Table A2 - VSS Document Capital Cost Breakdown for Ion Exchange Processes

Component Capital Cost 

Factor

Percent of Total 

Capital Cost

Capital Cost 

Breakdown 

Category

Equipment 1.0000 54.78% p

Installation 0.3000 16.43% c

Sitework/Interface Piping 0.0780 4.27% c

Standby Power 0.0650 3.56% c

OH&P 0.1732 9.49% e

Legal & Admin 0.0433 2.37% e

Engineering 0.1659 9.09% e

Contigencies 0.0000 0.00% c

Total 1.8254 100.00%

Table A3 - VSS Document Capital Cost Breakdown for Chlorination

Component Capital Cost 

Factor

Percent of Total 

Capital Cost

Capital Cost 

Breakdown 

Category

Equipment 1.0000 61.93% p

Installation 0.1500 9.29% c

Sitework/Interface Piping 0.0690 4.27% c

Standby Power 0.0575 3.56% c

OH&P 0.1532 9.49% e

Legal & Admin 0.0383 2.37% e

Engineering 0.1468 9.09% e

Contigencies 0.0000 0.00% c

Total 1.6148 100.00%

International Consultants, Inc.

Contract 68-C6-0039 A-1

February 1999 

VSS Document Capital Cost Breakdown 
Table A4 - VSS Document Capital Cost Breakdown for Potassium Permanganate Feed

Component Capital Cost 

Factor

Percent of Total 

Capital Cost

Capital Cost 

Breakdown 

Category

Equipment 1.0000 64.74% p

Installation 0.1000 6.47% c

Sitework/Interface Piping 0.0660 4.27% c

Standby Power 0.0550 3.56% c

OH&P 0.1465 9.49% e

Legal & Admin 0.0366 2.37% e

Engineering 0.1404 9.09% e

Contigencies 0.0000 0.00% c

Total 1.5446 100.00%

Table A5 - Typical VSS Document Capital Cost Breakdown

Component Capital Cost 

Factor

Percent of Total 

Capital Cost

Capital Cost 

Breakdown 

Category

Equipment 1.0000 54.78% p

Installation 0.3000 16.43% c

Sitework/Interface Piping 0.0780 4.27% c

Standby Power 0.0650 3.56% c

OH&P 0.1732 9.49% e

Legal & Admin 0.0433 2.37% e

Engineering 0.1659 9.09% e

Contigencies 0.0000 0.00% c

Total 1.8254 100.00%

International Consultants, Inc.

Contract 68-C6-0039 A-2

February 1999 

VSS Document Capital Cost Breakdown 
Appendix B Water Model Capital Cost

Breakdown Summaries

Table B1.1 - Base Costs Obtained from the Water Model for Activated Alumina

Contactor Volume (ft3)

32 71 126 283 385 502 754

Excavation & Sitework $4,700 $4,700 $4,700 $4,700 $4,700 $4,700 $4,700 c 

Manufactured Equipment $12,800 $23,900 $39,100 $50,600 $64,500 $72,900 $101,000 p

Activated Alumina $1,400 $3,100 $5,400 $11,900 $15,400 $19,600 $29,400 p

Concrete $400 $1,200 $1,800 $2,000 $2,500 $3,200 $4,100 p

Labor $1,200 $1,500 $2,000 $2,800 $3,300 $3,400 $4,200 c

Pipes and Valves $5,200 $6,500 $6,500 $8,400 $12,800 $13,300 $20,100 p

Electrical $6,400 $6,400 $6,400 $8,000 $8,000 $8,500 $9,600 p

Housing $8,700 $14,400 $16,900 $17,900 $24,800 $34,400 $43,900 p

Subtotal $40,800 $61,700 $82,800 $106,300 $136,000 $160,000 $217,000

Contingencies $6,100 $9,300 $12,400 $15,900 $20,400 $24,000 $32,600 c

Total $46,900 $71,000 $95,200 $122,200 $156,400 $184,000 $249,600

Table B1.2 - Water Model Base Construction Cost Analysis for Activated Alumina

Contactor Volume (ft3)

32 71 126 283 385 502 754

Excavation & Sitework 10.02% 6.62% 4.94% 3.85% 3.01% 2.55% 1.88% 4.70%

Manufactured Equipment 27.29% 33.66% 41.07% 41.41% 41.24% 39.62% 40.46% 37.82%

Activated Alumina 2.99% 4.37% 5.67% 9.74% 9.85% 10.65% 11.78% 7.86%

Concrete 0.85% 1.69% 1.89% 1.64% 1.60% 1.74% 1.64% 1.58%

Labor 2.56% 2.11% 2.10% 2.29% 2.11% 1.85% 1.68% 2.10%

Pipes and Valves 11.09% 9.15% 6.83% 6.87% 8.18% 7.23% 8.05% 8.20%

Electrical 13.65% 9.01% 6.72% 6.55% 5.12% 4.62% 3.85% 7.07%

Housing 18.55% 20.28% 17.75% 14.65% 15.86% 18.70% 17.59% 17.62%

Contingencies 13.01% 13.10% 13.03% 13.01% 13.04% 13.04% 13.06% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Capital Cost 

Category Cost Component

Cost Component

International Consultants, Inc.

Contract 68-C6-0039 B-1

February 1999 

Water Model Capital Cost Breakdown 
Table B2.1 - Base Costs Obtained from the Water Model for Anion Exchange

Resin Volume (ft3)

4 17 54 188 280 520

Excavation & Sitework $2,100 $2,100 $4,400 $4,400 $4,400 $5,300 c 

Manufactured Equipment $3,100 $8,600 $23,100 $64,100 $96,800 $164,800 p

Concrete $300 $400 $5,500 $5,800 $6,000 $8,400 p

Steel $0 $0 $7,800 $7,800 $7,800 $10,900 p

Labor $400 $1,100 $12,100 $12,800 $12,900 $17,200 c

Pipes and Valves $800 $800 $1,000 $2,600 $2,600 $3,100 p

Electrical $3,100 $3,100 $3,100 $3,100 $3,100 $3,100 p

Housing $5,600 $9,600 $11,100 $16,600 $19,200 $25,000 p

Subtotal $15,400 $25,700 $68,100 $117,200 $152,800 $237,800

Contingencies $2,300 $3,900 $10,200 $17,600 $22,900 $35,700 c

Total $17,700 $29,600 $78,300 $134,800 $175,700 $273,500

Table B2.2 - Water Model Base Construction Cost Analysis for Anion Exchange

Resin Volume (ft3)

4 17 54 188 280 520

Excavation & Sitework 11.86% 7.09% 5.62% 3.26% 2.50% 1.94% 5.38%

Manufactured Equipment 17.51% 29.05% 29.50% 47.55% 55.09% 60.26% 39.83%

Concrete 1.69% 1.35% 7.02% 4.30% 3.41% 3.07% 3.48%

Steel 0.00% 0.00% 9.96% 5.79% 4.44% 3.99% 4.03%

Labor 2.26% 3.72% 15.45% 9.50% 7.34% 6.29% 7.43%

Pipes and Valves 4.52% 2.70% 1.28% 1.93% 1.48% 1.13% 2.17%

Electrical 17.51% 10.47% 3.96% 2.30% 1.76% 1.13% 6.19%

Housing 31.64% 32.43% 14.18% 12.31% 10.93% 9.14% 18.44%

Contingencies 12.99% 13.18% 13.03% 13.06% 13.03% 13.05% 13.06%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Capital Cost 

Category Cost Component

Cost Component

International Consultants, Inc.

Contract 68-C6-0039 B-2

February 1999 

Water Model Capital Cost Breakdown 
Table B3.1 - Base Costs Obtained from the Water Model for Basic Chemical Feed

Maximum Feed Rate (lb/day)

0.1-10 25 50 100 250 500 1000

Dissolving Tank $290 $430 $640 $910 $1,830 $2,200 $4,400 p

Mixer $180 $200 $200 $240 $410 $620 $620 p

Metering Pump $430 $700 $750 $1,230 $1,600 $1,670 $1,820 p

Pipes and Valves $180 $180 $220 $220 $280 $280 $420 p

Labor $180 $180 $240 $260 $300 $330 $400 c

Electrical $80 $100 $150 $200 $250 $300 $400 p

Subtotal $1,340 $1,790 $2,200 $3,060 $4,670 $5,400 $8,060

Contingencies $200 $270 $330 $460 $700 $810 $1,210 c

Total $1,540 $2,060 $2,530 $3,520 $5,370 $6,210 $9,270

Table B3.2 - Water Model Base Construction Cost Analysis for Basic Chemical Feed

Maximum Feed Rate (lb/day)

0.1-10 25 50 100 250 500 1000

Dissolving Tank 18.83% 20.87% 25.30% 25.85% 34.08% 35.43% 47.46% 29.69%

Mixer 11.69% 9.71% 7.91% 6.82% 7.64% 9.98% 6.69% 8.63%

Metering Pump 27.92% 33.98% 29.64% 34.94% 29.80% 26.89% 19.63% 28.97%

Pipes and Valves 11.69% 8.74% 8.70% 6.25% 5.21% 4.51% 4.53% 7.09%

Labor 11.69% 8.74% 9.49% 7.39% 5.59% 5.31% 4.31% 7.50%

Electrical 5.19% 4.85% 5.93% 5.68% 4.66% 4.83% 4.31% 5.07%

Contingencies 12.99% 13.11% 13.04% 13.07% 13.04% 13.04% 13.05% 13.05%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Capital Cost 

Category Cost Component

Cost Component

International Consultants, Inc.

Contract 68-C6-0039 B-3

February 1999 

Water Model Capital Cost Breakdown 
Table B4.1 - Base Costs Obtained from the Water Model for Chlorination

Excavation & Sitework $1,200 c 

Manufactured Equipment $2,700 p

Concrete $300 p

Labor $400 c

Pipes and Valves $500 p

Electrical $2,200 p

Housing $7,800 p

Subtotal $15,100

Contingencies $2,300 c

Total $17,400

Table B4.2 - Water Model Base Construction Cost Analysis for Chlorination

Excavation & Sitework 6.90% 6.90%

Manufactured Equipment 15.52% 15.52%

Concrete 1.72% 1.72%

Labor 2.30% 2.30%

Pipes and Valves 2.87% 2.87%

Electrical 12.64% 12.64%

Housing 44.83% 44.83%

Contingencies 13.22% 13.22%

Total 100.00% 100.00%

Average

Percent

Cost 

Cost 

Capital Cost 

Category Cost Component

Cost Component

International Consultants, Inc.

Contract 68-C6-0039 B-4

February 1999 

Water Model Capital Cost Breakdown 
Table B5.1 - Base Costs Obtained from the Water Model for Underground Clearwell Storage

Design Capacity (gpd)

5,000 10,000 50,000 100,000 500,000

Excavation & Sitework $3,300 $5,700 $16,500 $25,300 $75,400 c 

Concrete $9,800 $16,500 $37,000 $64,000 $216,400 p

Steel $300 $400 $500 $500 $600 p

Electrical $2,600 $2,600 $2,600 $2,600 $2,600 p

Subtotal $16,000 $25,200 $56,600 $92,400 $295,000

Contingencies $2,400 $3,800 $8,500 $13,900 $44,300 c

Total $18,400 $29,000 $65,100 $106,300 $339,300

Table B5.2 - Water Model Base Construction Cost Analysis for Underground Clearwell Storage

Design Capacity (gpd)

5,000 10,000 50,000 100,000 500,000

Excavation & Sitework 17.93% 19.66% 25.35% 23.80% 22.22% 21.79%

Concrete 53.26% 56.90% 56.84% 60.21% 63.78% 58.20%

Steel 1.63% 1.38% 0.77% 0.47% 0.18% 0.88%

Electrical 14.13% 8.97% 3.99% 2.45% 0.77% 6.06%

Contingencies 13.04% 13.10% 13.06% 13.08% 13.06% 13.07%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Capital Cost 

Category

Average

Percent

Cost Component

Cost Component

International Consultants, Inc.

Contract 68-C6-0039 B-5

February 1999 

Water Model Capital Cost Breakdown 
Table B6.1 - Base Costs Obtained from the Water Model for Package Conventional Treatment

Filter Area (ft2)

2 12 20 40 112 150

Excavation & Sitework $3,500 $3,500 $4,700 $5,800 $7,000 $9,300 c 

Manufactured Equipment $31,000 $44,900 $53,500 $111,300 $176,600 $190,500 p

Concrete $1,000 $1,000 $1,500 $4,500 $5,700 $6,800 p

Labor $9,900 $14,700 $17,500 $36,400 $57,800 $62,400 c

Pipes and Valves $4,200 $8,300 $10,400 $20,900 $29,200 $41,700 p

Electrical $3,200 $4,500 $5,300 $11,100 $17,600 $19,000 p

Housing $18,600 $18,600 $23,400 $45,000 $47,500 $52,500 p

Subtotal $71,400 $95,500 $116,300 $235,000 $341,400 $382,200

Contingencies $10,700 $14,300 $17,400 $35,300 $51,200 $57,300 c

Total $82,100 $109,800 $133,700 $270,300 $392,600 $439,500

Table B6.2 - Water Model Base Construction Cost Analysis for Package Conventional Treatment

Filter Area (ft2)

2 12 20 40 112 150

Excavation & Sitework 4.26% 3.19% 3.52% 2.15% 1.78% 2.12% 2.84%

Manufactured Equipment 37.76% 40.89% 40.01% 41.18% 44.98% 43.34% 41.36%

Concrete 1.22% 0.91% 1.12% 1.66% 1.45% 1.55% 1.32%

Labor 12.06% 13.39% 13.09% 13.47% 14.72% 14.20% 13.49%

Pipes and Valves 5.12% 7.56% 7.78% 7.73% 7.44% 9.49% 7.52%

Electrical 3.90% 4.10% 3.96% 4.11% 4.48% 4.32% 4.15%

Housing 22.66% 16.94% 17.50% 16.65% 12.10% 11.95% 16.30%

Contingencies 13.03% 13.02% 13.01% 13.06% 13.04% 13.04% 13.03%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Capital Cost 

Category

Average

Percent

Cost Component

Cost Component

International Consultants, Inc.

Contract 68-C6-0039 B-6

February 1999 

Water Model Capital Cost Breakdown 
Table B7.1 - Base Costs Obtained from the Water Model for Ferric Chloride Feed

Maximum Feed Rate (lb/day)

1 10 25 50 100 250 750

Storage Tank $0 $0 $0 $0 $360 $780 $2,040 p

Wooden Stairway $0 $0 $0 $0 $0 $300 $300 p

Metering Pump $390 $390 $390 $390 $390 $1,090 $1,100 p

Pipes and Valves $180 $180 $180 $180 $220 $280 $280 p

Labor $120 $120 $130 $130 $210 $360 $410 c

Electrical $80 $80 $80 $80 $100 $120 $120 p

Subtotal $770 $770 $780 $780 $1,280 $2,930 $4,250

Contingencies $120 $120 $120 $120 $190 $440 $640 c

Total $890 $890 $900 $900 $1,470 $3,370 $4,890

Table B7.2 - Water Model Base Construction Cost Analysis for Ferric Chloride Feed

Maximum Feed Rate (lb/day)

1 10 25 50 100 250 750

Storage Tank 0.00% 0.00% 0.00% 0.00% 24.49% 23.15% 41.72% 12.76%

Wooden Stairway 0.00% 0.00% 0.00% 0.00% 0.00% 8.90% 6.13% 2.15%

Metering Pump 43.82% 43.82% 43.33% 43.33% 26.53% 32.34% 22.49% 36.53%

Pipes and Valves 20.22% 20.22% 20.00% 20.00% 14.97% 8.31% 5.73% 15.64%

Labor 13.48% 13.48% 14.44% 14.44% 14.29% 10.68% 8.38% 12.74%

Electrical 8.99% 8.99% 8.89% 8.89% 6.80% 3.56% 2.45% 6.94%

Contingencies 13.48% 13.48% 13.33% 13.33% 12.93% 13.06% 13.09% 13.24%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Capital Cost 

Category Cost Component

Cost Component

International Consultants, Inc.

Contract 68-C6-0039 B-7

February 1999 

Water Model Capital Cost Breakdown 
Table B8.1 - Base Costs Obtained from the Water Model for Package Lime Softening

Design Capacity (gpd)

15,000 150,000 430,000 750,000 1,000,000

Excavation & Sitework $3,500 $5,800 $6,700 $8,400 $9,800 c 

Manufactured Equipment $33,200 $49,800 $66,300 $86,200 $103,800 p

Concrete $1,100 $2,500 $3,200 $5,900 $7,000 p

Labor $14,000 $18,200 $28,000 $36,400 $43,800 c

Pipes and Valves $5,200 $10,400 $14,100 $16,700 $45,900 p

Electrical $8,500 $12,200 $17,000 $18,900 $26,700 p

Housing $8,800 $16,400 $19,800 $30,000 $33,000 p

Subtotal $74,300 $115,300 $155,100 $202,500 $270,000

Contingencies $11,100 $17,300 $23,300 $30,400 $40,500 c

Total $85,400 $132,600 $178,400 $232,900 $310,500

Table B8.2 - Water Model Base Construction Cost Analysis for Package Lime Softening

Design Capacity (gpd)

15,000 150,000 430,000 750,000 1,000,000

Excavation & Sitework 4.10% 4.37% 3.76% 3.61% 3.16% 3.80%

Manufactured Equipment 38.88% 37.56% 37.16% 37.01% 33.43% 36.81%

Concrete 1.29% 1.89% 1.79% 2.53% 2.25% 1.95%

Labor 16.39% 13.73% 15.70% 15.63% 14.11% 15.11%

Pipes and Valves 6.09% 7.84% 7.90% 7.17% 14.78% 8.76%

Electrical 9.95% 9.20% 9.53% 8.12% 8.60% 9.08%

Housing 10.30% 12.37% 11.10% 12.88% 10.63% 11.46%

Contingencies 13.00% 13.05% 13.06% 13.05% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Capital Cost 

Category Cost Component

Cost Component

International Consultants, Inc.

Contract 68-C6-0039 B-8

February 1999 

Water Model Capital Cost Breakdown 
Table B9.1 - Base Costs Obtained from the Water Model for Permanganate Feed

Maximum Feed Rate (lb/day)

0.5-5 12.5 25 50 125 250

Dissolving Tank $290 $430 $640 $910 $1,830 $2,200 p

Mixer $180 $200 $200 $240 $410 $620 p

Metering Pump $430 $700 $750 $1,230 $1,600 $1,670 p

Pipes and Valves $180 $180 $220 $220 $280 $280 p

Labor $180 $180 $240 $260 $300 $330 c

Electrical $80 $100 $150 $200 $250 $300 p

Subtotal $1,340 $1,790 $2,200 $3,060 $4,670 $5,400

Contingencies $200 $270 $330 $460 $700 $810 c

Total $1,540 $2,060 $2,530 $3,520 $5,370 $6,210

Table B9.2 - Water Model Base Construction Cost Analysis for Permanganate Feed

Maximum Feed Rate (lb/day)

0.5-5 12.5 25 50 125 250

Excavation & Sitework 18.83% 20.87% 25.30% 25.85% 34.08% 35.43% 26.73%

Manufactured Equipment 11.69% 9.71% 7.91% 6.82% 7.64% 9.98% 8.96%

Concrete 27.92% 33.98% 29.64% 34.94% 29.80% 26.89% 30.53%

Labor 11.69% 8.74% 8.70% 6.25% 5.21% 4.51% 7.52%

Pipes and Valves 11.69% 8.74% 9.49% 7.39% 5.59% 5.31% 8.03%

Electrical 5.19% 4.85% 5.93% 5.68% 4.66% 4.83% 5.19%

Contingencies 12.99% 13.11% 13.04% 13.07% 13.04% 13.04% 13.05%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Cost Component

Cost Component

Capital Cost 

Category

International Consultants, Inc.

Contract 68-C6-0039 B-9

February 1999 

Water Model Capital Cost Breakdown 
Table B10.1 - Base Costs Obtained from the Water Model for Polymer Feed

Maximum Feed Rate (lb/day)

0.6 1 2.1 4.2 10.4

Mixing Tank $290 $430 $640 $910 $1,830 p

Mixer $850 $850 $200 $1,050 $1,050 p

Metering Pump $640 $700 $750 $1,230 $1,600 p

Pipes and Valves $180 $180 $220 $220 $280 p

Labor $180 $180 $240 $260 $300 c

Electrical $80 $100 $150 $200 $250 p

Subtotal $2,220 $2,440 $2,200 $3,870 $5,310

Contingencies $330 $370 $330 $580 $800 c

Total $2,550 $2,810 $2,530 $4,450 $6,110

Table B10.2 - Water Model Base Construction Cost Analysis for Polymer Feed

Maximum Feed Rate (lb/day)

0.6 1 2.1 4.2 10.4

Mixing Tank 11.37% 15.30% 25.30% 20.45% 29.95% 20.47%

Mixer 33.33% 30.25% 7.91% 23.60% 17.18% 22.45%

Metering Pump 25.10% 24.91% 29.64% 27.64% 26.19% 26.70%

Pipes and Valves 7.06% 6.41% 8.70% 4.94% 4.58% 6.34%

Labor 7.06% 6.41% 9.49% 5.84% 4.91% 6.74%

Electrical 3.14% 3.56% 5.93% 4.49% 4.09% 4.24%

Contingencies 12.94% 13.17% 13.04% 13.03% 13.09% 13.06%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Cost Component

Cost Component

Capital Cost 

Category

International Consultants, Inc.

Contract 68-C6-0039 B-10

February 1999 

Water Model Capital Cost Breakdown 
Table B11.1 - Base Costs Obtained from the Water Model for Raw Water Pumping

Design Capacity (gpd)

28,800 144,000 504,000 720,000 1,008,000

Excavation & Sitework $11,700 $11,700 $12,300 $12,300 $12,800 c 

Manufactured Equipment $6,600 $7,800 $11,800 $12,600 $16,500 p

Concrete $500 $500 $1,100 $1,100 $1,500 p

Labor $3,700 $3,800 $5,800 $6,200 $8,500 c

Pipes and Valves $1,500 $1,800 $2,700 $3,600 $4,500 p

Electrical $800 $800 $1,400 $1,600 $2,100 p

Subtotal $24,800 $26,400 $35,100 $37,400 $45,900

Contingencies $3,700 $4,000 $5,300 $5,600 $6,900 c

Total $28,500 $30,400 $40,400 $43,000 $52,800

Table B11.2 - Water Model Base Construction Cost Analysis for Raw Water Pumping

Design Capacity (gpd)

28,800 144,000 504,000 720,000 1,008,000

Excavation & Sitework 41.05% 38.49% 30.45% 28.60% 24.24% 32.57%

Manufactured Equipment 23.16% 25.66% 29.21% 29.30% 31.25% 27.72%

Concrete 1.75% 1.64% 2.72% 2.56% 2.84% 2.30%

Labor 12.98% 12.50% 14.36% 14.42% 16.10% 14.07%

Pipes and Valves 5.26% 5.92% 6.68% 8.37% 8.52% 6.95%

Electrical 2.81% 2.63% 3.47% 3.72% 3.98% 3.32%

Contingencies 12.98% 13.16% 13.12% 13.02% 13.07% 13.07%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Cost Component

Cost Component

Capital Cost 

Category

International Consultants, Inc.

Contract 68-C6-0039 B-11

February 1999 

Water Model Capital Cost Breakdown 
Table B12.1 - Base Costs Obtained from the Water Model for Package Reverse Osmosis

Plant Capacity (gpd)

2,500 10,000 50,000 100,000 500,000 1,000,000

Manufactured Equipment $20,300 $30,000 $69,600 $123,000 $454,800 $877,400 p

Labor $800 $1,200 $1,500 $2,800 $7,500 $14,600 c

Electrical $3,200 $4,600 $10,700 $18,700 $45,900 $62,100 p

Housing $11,900 $13,900 $16,400 $18,500 $38,400 $52,500 p

Subtotal $36,200 $49,700 $98,200 $163,000 $546,600 $1,006,600

Contingencies $5,400 $7,500 $14,700 $24,500 $82,000 $151,000 c

Total $41,600 $57,200 $112,900 $187,500 $628,600 $1,157,600

Table B12.2 - Water Model Base Construction Cost Analysis for Package Reverse Osmosis

Plant Capacity (gpd)

2,500 10,000 50,000 100,000 500,000 1,000,000

Manufactured Equipment 48.80% 52.45% 61.65% 65.60% 72.35% 75.79% 62.77%

Labor 1.92% 2.10% 1.33% 1.49% 1.19% 1.26% 1.55%

Electrical 7.69% 8.04% 9.48% 9.97% 7.30% 5.36% 7.98%

Housing 28.61% 24.30% 14.53% 9.87% 6.11% 4.54% 14.66%

Contingencies 12.98% 13.11% 13.02% 13.07% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Cost Component

Cost Component

Capital Cost 

Category

International Consultants, Inc.

Contract 68-C6-0039 B-12

February 1999 

Water Model Capital Cost Breakdown 
Table B13.1 - Base Costs Obtained from the Water Model for Sodium Hydroxide Feed

Maximum Feed Rate (lb/day)

0.8 4 8 42 83 417 834

Storage and Feed Tanks $60 $60 $90 $970 $2,040 $3,560 $6,940 p

Heating and Insulation $0 $0 $0 $200 $410 $950 $1,620 p

Mixer $0 $0 $0 $180 $240 $620 $640 p

Stairway $0 $0 $0 $0 $0 $300 $600 p

Man. Transfer Pump $100 $100 $100 $0 $0 $0 $0 p

Pipes and Valves $310 $310 $310 $470 $470 $530 $790 p

Metering Pump $390 $390 $390 $390 $410 $1,090 $1,100 p

Containment Wall $120 $120 $150 $270 $400 $600 $880 p

Labor $280 $280 $280 $420 $480 $650 $860 c

Electrical $80 $80 $80 $100 $100 $120 $120 p

Subtotal $1,340 $1,340 $1,400 $3,000 $4,550 $8,420 $13,550

Contingencies $200 $200 $210 $450 $680 $1,260 $2,030 c

Total $1,540 $1,540 $1,610 $3,450 $5,230 $9,680 $15,580

Table B13.2 - Water Model Base Construction Cost Analysis for Sodium Hydroxide Feed

Maximum Feed Rate (lb/day)

0.8 4 8 42 83 417 834

Storage and Feed Tanks 3.90% 3.90% 5.59% 28.12% 39.01% 36.78% 44.54% 23.12%

Heating and Insulation 0.00% 0.00% 0.00% 5.80% 7.84% 9.81% 10.40% 4.84%

Mixer 0.00% 0.00% 0.00% 5.22% 4.59% 6.40% 4.11% 2.90%

Stairway 0.00% 0.00% 0.00% 0.00% 0.00% 3.10% 3.85% 0.99%

Man. Transfer Pump 6.49% 6.49% 6.21% 0.00% 0.00% 0.00% 0.00% 2.74%

Pipes and Valves 20.13% 20.13% 19.25% 13.62% 8.99% 5.48% 5.07% 13.24%

Metering Pump 25.32% 25.32% 24.22% 11.30% 7.84% 11.26% 7.06% 16.05%

Containment Wall 7.79% 7.79% 9.32% 7.83% 7.65% 6.20% 5.65% 7.46%

Labor 18.18% 18.18% 17.39% 12.17% 9.18% 6.71% 5.52% 12.48%

Electrical 5.19% 5.19% 4.97% 2.90% 1.91% 1.24% 0.77% 3.17%

Contingencies 12.99% 12.99% 13.04% 13.04% 13.00% 13.02% 13.03% 13.02%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Cost Component

Cost Component

Capital Cost 

Category

International Consultants, Inc.

Contract 68-C6-0039 B-13

February 1999 

Water Model Capital Cost Breakdown 
Table B14.1 - Base Costs Obtained from the Water Model for Package Ultrafiltration

Membrane Area (ft2)

30 424 1,431 3,604 7,155 14,310

Excavation & Sitework $1,300 $2,400 $4,100 $5,700 $10,200 $14,900 c

Manufactured Equipment $5,500 $25,300 $65,600 $129,800 $23,900 $415,100 p

Concrete $1,800 $3,700 $5,800 $10,200 $16,700 $28,800 p

Labor $1,100 $5,200 $13,500 $26,900 $49,500 $85,900 c

Pipes and Valves $500 $1,100 $2,200 $3,800 $4,500 $6,200 p

Electrical $1,500 $5,600 $13,300 $25,800 $48,100 $85,300 p

Housing $7,800 $14,600 $21,700 $29,000 $40,800 $56,000 p

Subtotal $19,500 $57,900 $126,200 $231,200 $193,700 $692,200

Contingencies $2,900 $8,700 $18,900 $34,700 $29,100 $103,800 c

Total $22,400 $66,600 $145,100 $265,900 $222,800 $796,000

Table B14.2 - Water Model Base Construction Cost Analysis for Package Ultrafiltration

Membrane Area (ft2)

30 424 1,431 3,604 7,155 14,310

Excavation & Sitework 5.80% 3.60% 2.83% 2.14% 4.58% 1.87% 3.47%

Manufactured Equipment 24.55% 37.99% 45.21% 48.82% 10.73% 52.15% 36.57%

Concrete 8.04% 5.56% 4.00% 3.84% 7.50% 3.62% 5.42%

Labor 4.91% 7.81% 9.30% 10.12% 22.22% 10.79% 10.86%

Pipes and Valves 2.23% 1.65% 1.52% 1.43% 2.02% 0.78% 1.60%

Electrical 6.70% 8.41% 9.17% 9.70% 21.59% 10.72% 11.05%

Housing 34.82% 21.92% 14.96% 10.91% 18.31% 7.04% 17.99%

Contingencies 12.95% 13.06% 13.03% 13.05% 13.06% 13.04% 13.03%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Average

Percent

Cost Component

Cost Component

Capital Cost 

Category

International Consultants, Inc.

Contract 68-C6-0039 B-14

February 1999 

Water Model Capital Cost Breakdown 
Appendix C W/W Cost Model Capital Cost

Breakdown Summaries

Table C1.1 - Base Costs Obtained from the WATERCO$T Model for Activated Alumina

Plant Capacity (mgd)

0.7 2.0 6.8 27 54 135

Manufactured Equipment $26,760 $44,580 $138,330 $522,210 $1,031,270 $2,564,560 p

Activated Alumina $8,300 $14,770 $83,080 $332,310 $664,610 $1,661,530 p

Labor $10,280 $13,490 $48,010 $192,020 $384,060 $1,282,370 c 

Pipes and Valves $16,260 $19,320 $69,030 $273,210 $542,650 $1,368,060 p

Electrical $10,050 $11,360 $22,300 $60,300 $119,030 $284,750 p

Housing $6,960 $27,630 $62,120 $210,980 $374,840 $744,320 p

Contingencies $11,790 $19,670 $63,430 $238,650 $467,470 $1,185,840 c

Total $90,400 $150,820 $486,300 $1,829,680 $3,583,930 $9,091,430

Table C1.2 - WATERCO$T Model Base Construction Cost Analysis for Activated Alumina

Plant Capacity (mgd)

0.7 2.0 6.8 27 54 135

Manufactured Equipment 29.60% 29.56% 28.45% 28.54% 28.77% 28.21% 28.86%

Activated Alumina 9.18% 9.79% 17.08% 18.16% 18.54% 18.28% 15.17%

Labor 11.37% 8.94% 9.87% 10.49% 10.72% 14.11% 10.92%

Pipes and Valves 17.99% 12.81% 14.19% 14.93% 15.14% 15.05% 15.02%

Electrical 11.12% 7.53% 4.59% 3.30% 3.32% 3.13% 5.50%

Housing 7.70% 18.32% 12.77% 11.53% 10.46% 8.19% 11.49%

Contingencies 13.04% 13.04% 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-1

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C2.1 - Base Costs Obtained from the WATERCO$T Model for Ammonia Feed Systems

Feed Capacity (lb/day)

250 500 1,000 2,500 5,000

Manufactured Equipment $13,260 $19,520 $30,450 $38,830 $59,200 p

Labor $3,990 $5,680 $9,250 $10,620 $13,870 c

Pipes and Valves $2,390 $3,520 $5,500 $7,000 $10,670 p

Electrical $3,250 $3,770 $6,180 $8,480 $10,990 p

Housing $4,500 $4,500 $4,500 $4,500 $6,430 p

Contingencies $4,110 $5,550 $8,380 $10,410 $15,170 c

Total $31,500 $42,540 $64,260 $79,840 $116,330

Table C2.2 - WATERCO$T Model Base Construction Cost Analysis for Ammonia Feed Systems

Feed Capacity (lb/day)

250 500 1,000 2,500 5,000

Manufactured Equipment 42.10% 45.89% 47.39% 48.63% 50.89% 46.98%

Labor 12.67% 13.35% 14.39% 13.30% 11.92% 13.13%

Pipes and Valves 7.59% 8.27% 8.56% 8.77% 9.17% 8.47%

Electrical 10.32% 8.86% 9.62% 10.62% 9.45% 9.77%

Housing 14.29% 10.58% 7.00% 5.64% 5.53% 8.61%

Contingencies 13.05% 13.05% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-2

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C3.1 - Base Costs Obtained from the WATERCO$T Model for Backwash Water Pumping

Pumping Capacity (mgd(gpm))

1,260

(1.8)

3,150

(4.5)

6,300

(9.1)

18,000

(25.9)

22,950

(33)

Manufactured Equipment $11,400 $14,600 $38,380 $76,780 $95,970 p

Labor $3,050 $4,410 $4,880 $9,290 $12,440 c

Pipes and Valves $9,780 $17,690 $17,690 $33,390 $44,780 p

Electrical $13,350 $16,040 $16,740 $28,070 $33,250 p

Contingencies $5,640 $7,910 $11,650 $22,130 $27,970 c

Total $43,220 $60,650 $89,340 $169,660 $214,410

Table C3.2 - WATERCO$T Model Base Construction Cost Analysis for Backwash Water Pumping

Pumping Capacity (mgd(gpm))

1,260

(1.8)

3,150

(4.5)

6,300

(9.1)

18,000

(25.9)

22,950

(33)

Manufactured Equipment 26.38% 24.07% 42.96% 45.26% 44.76% 36.68%

Labor 7.06% 7.27% 5.46% 5.48% 5.80% 6.21%

Pipes and Valves 22.63% 29.17% 19.80% 19.68% 20.89% 22.43%

Electrical 30.89% 26.45% 18.74% 16.54% 15.51% 21.63%

Contingencies 13.05% 13.04% 13.04% 13.04% 13.05% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-3

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C4.1 - Base Costs Obtained from the WATERCO$T Model for Chemical Sludge Pumping

Capacity (gpm)

20 100 500 1,000 5,000 10,000

Excavation & Sitework $470 $600 $810 $970 $1,840 $2,220 c

Manufactured Equipment $4,370 $6,230 $8,210 $10,390 $23,320 $38,440 p

Concrete $1,500 $2,210 $3,220 $4,100 $9,270 $12,310 p

Steel $1,510 $2,130 $3,120 $3,940 $8,640 $11,070 p

Labor $5,280 $8,060 $12,880 $17,400 $47,850 $64,720 c

Pipes and Valves $2,560 $4,570 $10,870 $18,190 $42,810 $79,060 p

Electrical $6,290 $7,390 $7,880 $9,380 $10,380 $12,510 p

Housing $5,880 $5,880 $5,880 $8,100 $8,100 $11,700 p

Contingencies $4,180 $5,560 $7,930 $10,870 $22,830 $34,800 c

Total $32,040 $42,630 $60,800 $83,340 $175,040 $266,830

Table C4.2 - WATERCO$T Model Base Construction Cost Analysis for Chemical Sludge Pumping

Capacity (gpm)

20 100 500 1,000 5,000 10,000

Excavation & Sitework 1.47% 1.41% 1.33% 1.16% 1.05% 0.83% 1.21%

Manufactured Equipment 13.64% 14.61% 13.50% 12.47% 13.32% 14.41% 13.66%

Concrete 4.68% 5.18% 5.30% 4.92% 5.30% 4.61% 5.00%

Steel 4.71% 5.00% 5.13% 4.73% 4.94% 4.15% 4.78%

Labor 16.48% 18.91% 21.18% 20.88% 27.34% 24.26% 21.51%

Pipes and Valves 7.99% 10.72% 17.88% 21.83% 24.46% 29.63% 18.75%

Electrical 19.63% 17.34% 12.96% 11.26% 5.93% 4.69% 11.97%

Housing 18.35% 13.79% 9.67% 9.72% 4.63% 4.38% 10.09%

Contingencies 13.05% 13.04% 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-4

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C5.1 - Base Costs Obtained from the WATERCO$T Model for Chlorination

Chlorine Feed Capacity (lb/day)

10 500 1,000 2,000 5,000 10,000

Manufactured Equipment $6,760 $21,630 $41,630 $65,950 $76,780 $114,360 p

Labor $820 $2,610 $5,030 $7,960 $9,270 $13,810 c

Pipes and Valves $540 $1,710 $3,300 $5,230 $6,080 $9,060 p

Electrical $770 $2,450 $4,710 $7,460 $8,690 $12,940 p

Housing $2,430 $18,360 $27,760 $46,550 $100,440 $186,490 p

Contingencies $1,700 $7,010 $12,360 $19,970 $30,190 $50,500 c

Total $13,020 $53,770 $94,790 $153,120 $231,450 $387,160

Table C5.2 - WATERCO$T Model Base Construction Cost Analysis for Chlorination

Chlorine Feed Capacity (lb/day)

10 500 1,000 2,000 5,000 10,000

Manufactured Equipment 51.92% 40.23% 43.92% 43.07% 33.17% 29.54% 40.31%

Labor 6.30% 4.85% 5.31% 5.20% 4.01% 3.57% 4.87%

Pipes and Valves 4.15% 3.18% 3.48% 3.42% 2.63% 2.34% 3.20%

Electrical 5.91% 4.56% 4.97% 4.87% 3.75% 3.34% 4.57%

Housing 18.66% 34.15% 29.29% 30.40% 43.40% 48.17% 34.01%

Contingencies 13.06% 13.04% 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-5

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C6.1 - Base Costs Obtained from the WATERCO$T Model for Circular Clarifiers

Surface Area (SA=ft2) and Diameter (D=ft)

SA=707

 D=30

SA=1,590

D=45

SA=5,027

D=80

SA=10,387

D=115

SA=15,393

D=140

SA=22,698

D=170

SA=31,416

D=200

Excavation & Sitework $1,530 $2,430 $4,900 $7,860 $10,280 $13,520 $17,130 c

Manufactured Equipment $28,740 $34,410 $69,580 $97,180 $132,350 $189,060 $226,980 p

Concrete $4,860 $7,710 $15,480 $24,800 $32,400 $42,560 $53,860 p

Steel $14,160 $21,090 $67,240 $129,250 $188,720 $249,570 $335,140 p

Labor $10,770 $16,180 $30,960 $46,980 $60,110 $77,640 $96,320 c

Pipes and Valves $8,090 $8,420 $11,540 $15,660 $21,590 $26,590 $42,520 p

Electrical $5,940 $5,940 $7,560 $8,270 $10,870 $12,370 $13,060 p

Contingencies $11,110 $14,430 $31,090 $49,500 $68,450 $91,700 $117,750 c

Total $85,200 $110,610 $238,350 $379,500 $524,770 $703,010 $902,760

Table C6.2 - WATERCO$T Model Base Construction Cost Analysis for Circular Clarifiers

Surface Area (SA=ft2) and Diameter (D=ft)

SA=707

 D=30

SA=1,590

D=45

SA=5,027

D=80

SA=10,387

D=115

SA=15,393

D=140

SA=22,698

D=170

SA=31,416

D=200

Excavation & Sitework 1.80% 2.20% 2.06% 2.07% 1.96% 1.92% 1.90% 1.99%

Manufactured Equipment 33.73% 31.11% 29.19% 25.61% 25.22% 26.89% 25.14% 28.13%

Concrete 5.70% 6.97% 6.49% 6.53% 6.17% 6.05% 5.97% 6.27%

Steel 16.62% 19.07% 28.21% 34.06% 35.96% 35.50% 37.12% 29.51%

Labor 12.64% 14.63% 12.99% 12.38% 11.45% 11.04% 10.67% 12.26%

Pipes and Valves 9.50% 7.61% 4.84% 4.13% 4.11% 3.78% 4.71% 5.53%

Electrical 6.97% 5.37% 3.17% 2.18% 2.07% 1.76% 1.45% 3.28%

Contingencies 13.04% 13.05% 13.04% 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-6

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C7.1 - Base Costs Obtained from the WATERCO$T Model for Clearwell Storage

Capacity (gal)

10,000 50,000 100,000 500,000 1,000,000 7,500,000

Excavation & Sitework $140 $190 $410 $2,030 $19,440 $30,020 c

Concrete $8,250 $14,430 $23,280 $66,330 $105,520 $622,500 p

Steel $5,700 $9,240 $14,550 $32,670 $113,050 $350,700 p

Labor $13,050 $21,480 $35,040 $84,090 $109,290 $394,160 p

Electrical $1,270 $1,270 $6,010 $6,010 $9,800 $9,800 p

Contingencies $4,260 $6,990 $11,890 $28,670 $53,570 $211,080 c

Total $32,670 $53,600 $91,180 $219,800 $410,670 $1,618,260

Table C7.2 - WATERCO$T Model Base Construction Cost Analysis for Clearwell Storage

Capacity (gal)

10,000 50,000 100,000 500,000 1,000,000 7,500,000

Excavation & Sitework 0.43% 0.35% 0.45% 0.92% 4.73% 1.86% 1.46%

Concrete 25.25% 26.92% 25.53% 30.18% 25.69% 38.47% 28.67%

Steel 17.45% 17.24% 15.96% 14.86% 27.53% 21.67% 19.12%

Labor 39.94% 40.07% 38.43% 38.26% 26.61% 24.36% 34.61%

Electrical 3.89% 2.37% 6.59% 2.73% 2.39% 0.61% 3.10%

Contingencies 13.04% 13.04% 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-7

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C8.1 - Base Costs Obtained from the WATERCO$T Model for Ferric Chloride Feed Systems*

Feed Capacity (lb/hr)

10.7 107 1,070 5,350

Manufactured Equipment $7,500 $13,100 $33,560 $160,940 p

Labor $420 $1,130 $2,430 $12,160 c

Pipes and Valves $2,000 $2,500 $3,000 $15,000 p

Electrical $1,110 $2,260 $4,960 $19,000 p

Housing $6,000 $13,300 $51,270 $174,590 p

Contingencies $2,550 $4,840 $14,280 $57,250 c

Total $19,580 $37,130 $109,500 $438,940

*Numbers were unavailable for ferric chloride. However, numbers presented for ferrous sulfate and ferric sulfate were identical.

It was assumed that these same relationships apply to ferric chloride

Table C8.2 - WATERCO$T Model Base Construction Cost Analysis for Ferric Chloride Feed Systems*

Feed Capacity (lb/hr)

10.7 107 1,070 5,350

Manufactured Equipment 38.30% 35.28% 30.65% 36.67% 35.22%

Labor 2.15% 3.04% 2.22% 2.77% 2.54%

Pipes and Valves 10.21% 6.73% 2.74% 3.42% 5.78%

Electrical 5.67% 6.09% 4.53% 4.33% 5.15%

Housing 30.64% 35.82% 46.82% 39.78% 38.27%

Contingencies 13.02% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00%

*Numbers were unavailable for ferric chloride. However, numbers presented for ferrous sulfate and ferric sulfate were identical.

It was assumed that these same relationships apply to ferric chloride

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-8

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C9.1 - Base Costs Obtained from the WATERCO$T Model for Finished Water Pumping

Plant Capacity (mgd)

1.5 15 150 300

Manufactured Equipment $15,410 $89,700 $567,600 $1,142,350 p

Labor $3,880 $11,580 $80,400 $158,840 c

Pipes and Valves $5,200 $16,570 $139,200 $270,100 p

Electrical $7,180 $38,450 $210,490 $400,230 p

Contingencies $4,750 $23,450 $149,650 $295,730 c

Total $36,420 $179,750 $1,147,340 $2,267,250

Table C9.2 - WATERCO$T Model Base Construction Cost Analysis for Finished Water Pumping

Plant Capacity (mgd)

1.5 15 150 300

Manufactured Equipment 42.31% 49.90% 49.47% 50.38% 48.02%

Labor 10.65% 6.44% 7.01% 7.01% 7.78%

Pipes and Valves 14.28% 9.22% 12.13% 11.91% 11.89%

Electrical 19.71% 21.39% 18.35% 17.65% 19.28%

Contingencies 13.04% 13.05% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-9

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C10.1 - Base Costs Obtained from the WATERCO$T Model for Gravity Filtration

Total Filter Area (FA-ft2) and Plant Flow (Q=mgd)

FA=140

Q=1

FA=700

Q=5

FA=1,400

Q=10

FA=7,000

Q=50

FA=14,000

Q=100

FA=28,000

Q=200

Excavation & Sitework $1,950 $3,620 $5,520 $16,220 $25,590 $43,410 c

Manufactured Equipment $26,360 $56,960 $78,300 $305,170 $529,360 $982,390 p

Concrete $13,400 $27,040 $41,660 $95,490 $154,790 $275,570 p

Steel $11,550 $19,960 $30,120 $73,530 $123,160 $209,960 p

Labor $40,580 $88,490 $150,870 $356,380 $508,980 $1,000,670 c

Pipes and Valves $20,580 $79,020 $127,340 $420,670 $590,150 $1,125,500 p

Electrical $13,390 $38,410 $38,410 $99,140 $168,840 $265,310 p

Housing $17,400 $40,480 $70,590 $291,940 $514,330 $968,520 p

Contingencies $21,780 $53,100 $81,420 $248,780 $392,280 $730,700 c

Total $166,990 $407,080 $624,230 $1,907,320 $3,007,480 $5,602,030

Table C10.2 - WATERCO$T Model Base Construction Cost Analysis for Gravity Filtration

Total Filter Area (FA-ft2) and Plant Flow (Q=mgd)

FA=140

Q=1

FA=700

Q=5

FA=1,400

Q=10

FA=7,000

Q=50

FA=14,000

Q=100

FA=28,000

Q=200

Excavation & Sitework 1.17% 0.89% 0.88% 0.85% 0.85% 0.77% 0.90%

Manufactured Equipment 15.79% 13.99% 12.54% 16.00% 17.60% 17.54% 15.58%

Concrete 8.02% 6.64% 6.67% 5.01% 5.15% 4.92% 6.07%

Steel 6.92% 4.90% 4.83% 3.86% 4.10% 3.75% 4.72%

Labor 24.30% 21.74% 24.17% 18.68% 16.92% 17.86% 20.61%

Pipes and Valves 12.32% 19.41% 20.40% 22.06% 19.62% 20.09% 18.98%

Electrical 8.02% 9.44% 6.15% 5.20% 5.61% 4.74% 6.53%

Housing 10.42% 9.94% 11.31% 15.31% 17.10% 17.29% 13.56%

Contingencies 13.04% 13.04% 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-10

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C11.1 - Base Costs Obtained from the WATERCO$T Model for Horizontal Paddle, G=50

Total Basin Volume (ft3)

1,800 10,000 25,000 100,000 500,000 1,000,000

Excavation & Sitework $470 $2,550 $4,290 $9,970 $40,080 $77,640 p

Manufactured Equipment $12,140 $28,250 $35,410 $74,400 $220,800 $433,640 p

Concrete $1,400 $7,610 $12,740 $29,770 $120,280 $232,960 p

Steel $2,360 $12,550 $20,440 $46,500 $175,290 $339,510 p

Labor $7,080 $20,220 $29,420 $75,460 $221,200 $439,770 c

Electrical $6,980 $28,320 $28,320 $28,320 $141,610 $283,220 p

Contingencies $4,560 $14,930 $19,590 $39,660 $137,890 $271,010 c

Total $34,990 $114,430 $150,210 $304,080 $1,057,150 $2,077,750

Table C11.2 - WATERCO$T Model Base Construction Cost Analysis for Horizontal Paddle, G=50

Total Basin Volume (ft3)

1,800 10,000 25,000 100,000 500,000 1,000,000

Excavation & Sitework 1.34% 2.23% 2.86% 3.28% 3.79% 3.74% 2.87%

Manufactured Equipment 34.70% 24.69% 23.57% 24.47% 20.89% 20.87% 24.86%

Concrete 4.00% 6.65% 8.48% 9.79% 11.38% 11.21% 8.59%

Steel 6.74% 10.97% 13.61% 15.29% 16.58% 16.34% 13.26%

Labor 20.23% 17.67% 19.59% 24.82% 20.92% 21.17% 20.73%

Electrical 19.95% 24.75% 18.85% 9.31% 13.40% 13.63% 16.65%

Contingencies 13.03% 13.05% 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-11

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C12.1 - Base Costs Obtained from the WATERCO$T Model for Horizontal Paddle, G=80

Total Basin Volume (ft3)

1,800 10,000 25,000 100,000 500,000

Excavation & Sitework $470 $2,550 $4,290 $9,970 $40,080 c

Manufactured Equipment $12,140 $34,210 $44,360 $115,770 $427,670 p

Concrete $1,400 $7,610 $12,740 $29,770 $120,280 p

Steel $2,360 $12,550 $20,440 $46,500 $175,290 p

Labor $7,080 $22,190 $32,370 $90,170 $289,520 p

Electrical $6,980 $28,320 $28,320 $28,320 $141,610 p

Contingencies $4,560 $16,110 $21,380 $48,080 $179,170 c

Total $34,990 $123,540 $163,900 $368,580 $1,373,620

Table C12.2 - WATERCO$T Model Base Construction Cost Analysis for Horizontal Paddle, G=80

Total Basin Volume (ft3)

1,800 10,000 25,000 100,000 500,000

Excavation & Sitework 1.34% 2.06% 2.62% 2.70% 2.92% 2.33%

Manufactured Equipment 34.70% 27.69% 27.07% 31.41% 31.13% 30.40%

Concrete 4.00% 6.16% 7.77% 8.08% 8.76% 6.95%

Steel 6.74% 10.16% 12.47% 12.62% 12.76% 10.95%

Labor 20.23% 17.96% 19.75% 24.46% 21.08% 20.70%

Electrical 19.95% 22.92% 17.28% 7.68% 10.31% 15.63%

Contingencies 13.03% 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-12

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C13.1 - Base Costs Obtained from the WATERCO$T Model for Hydraulic Surface Wash

Total Filter Area (ft2)

140 700 1,400 7,000 14,000 28,000

Manufactured Equipment $9,170 $12,050 $35,090 $82,010 $172,440 $401,200 p

Labor $1,300 $2,770 $5,170 $14,710 $29,430 $66,600 c

Pipes and Valves $2,570 $5,100 $7,020 $13,390 $32,290 $59,870 p

Electrical $12,670 $17,920 $20,440 $37,900 $61,120 $92,360 p

Contingencies $3,860 $5,680 $10,160 $22,200 $44,290 $93,000 c

Total $29,570 $43,520 $77,880 $170,210 $339,570 $713,030

Table C13.2 - WATERCO$T Model Base Construction Cost Analysis for Hydraulic Surface Wash

Total Filter Area (ft2)

140 700 1,400 7,000 14,000 28,000

Manufactured Equipment 31.01% 27.69% 45.06% 48.18% 50.78% 56.27% 43.16%

Labor 4.40% 6.36% 6.64% 8.64% 8.67% 9.34% 7.34%

Pipes and Valves 8.69% 11.72% 9.01% 7.87% 9.51% 8.40% 9.20%

Electrical 42.85% 41.18% 26.25% 22.27% 18.00% 12.95% 27.25%

Contingencies 13.05% 13.05% 13.05% 13.04% 13.04% 13.04% 13.05%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component

Capital Cost

Category

Average 

Percent

International Consultants, Inc.

Contract 68-C6-0039 C-13

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C14.1 - Base Costs Obtained from the WATERCO$T Model for In-Plant Pumping

Pumping Capacity (mgd)

1 5 10 50 100 200

Excavation & Sitework $100 $100 $130 $360 $600 $1,030 c

Manufactured Equipment $6,300 $9,110 $14,780 $48,650 $83,400 $152,900 p

Concrete $970 $970 $1,510 $4,770 $8,030 $14,090 p

Steel $1,610 $1,610 $2,450 $7,630 $12,500 $21,330 p

Labor $5,570 $10,410 $24,070 $63,330 $129,130 $331,030 c

Pipes and Valves $5,090 $12,330 $16,300 $60,230 $114,200 $222,080 p

Electrical $3,170 $4,930 $7,390 $25,760 $47,240 $89,360 p

Housing $1,500 $1,500 $3,000 $14,520 $28,830 $58,080 p

Contingencies $3,650 $6,140 $10,440 $33,790 $63,590 $133,490 c

Total $27,960 $47,100 $80,070 $259,040 $487,520 $1,023,390

Table C14.2 - WATERCO$T Model Base Construction Cost Analysis for In-Plant Pumping

Pumping Capacity (mgd)

1 5 10 50 100 200

Excavation & Sitework 0.36% 0.21% 0.16% 0.14% 0.12% 0.10% 0.18%

Manufactured Equipment 22.53% 19.34% 18.46% 18.78% 17.11% 14.94% 18.53%

Concrete 3.47% 2.06% 1.89% 1.84% 1.65% 1.38% 2.05%

Steel 5.76% 3.42% 3.06% 2.95% 2.56% 2.08% 3.31%

Labor 19.92% 22.10% 30.06% 24.45% 26.49% 32.35% 25.89%

Pipes and Valves 18.20% 26.18% 20.36% 23.25% 23.42% 21.70% 22.19%

Electrical 11.34% 10.47% 9.23% 9.94% 9.69% 8.73% 9.90%

Housing 5.36% 3.18% 3.75% 5.61% 5.91% 5.68% 4.92%

Contingencies 13.05% 13.04% 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component

Capital Cost

Category

Average 

Percent

International Consultants, Inc.

Contract 68-C6-0039 C-14

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C15.1 - Base Costs Obtained from the WATERCO$T Model for Ion Exchange

Plant Capacity (mgd)

1.1 3.7 6.1 12.3

Excavation & Sitework $740 $1,140 $1,470 $1,970 c

Manufactured Equipment $39,960 $89,580 $137,770 $258,230 p

Media $92,790 $313,160 $521,940 $1,043,880 p

Concrete $2,410 $3,580 $4,750 $6,320 p

Steel $3,830 $5,680 $7,530 $9,950 p

Labor $17,420 $33,510 $61,460 $125,080 c

Pipes and Valves $14,040 $38,780 $69,740 $139,480 p

Electrical $27,700 $38,510 $60,820 $120,210 p

Housing $21,920 $35,660 $57,440 $79,820 p

Contingencies $33,120 $83,940 $138,440 $267,740 c

Total $253,930 $643,540 $1,061,360 $2,052,680

Table C15.2 - WATERCO$T Model Base Construction Cost Analysis for Ion Exchange

Plant Capacity (mgd)

1.1 3.7 6.1 12.3

Excavation & Sitework 0.29% 0.18% 0.14% 0.10% 0.18%

Manufactured Equipment 15.74% 13.92% 12.98% 12.58% 13.80%

Media 36.54% 48.66% 49.18% 50.85% 46.31%

Concrete 0.95% 0.56% 0.45% 0.31% 0.57%

Steel 1.51% 0.88% 0.71% 0.48% 0.90%

Labor 6.86% 5.21% 5.79% 6.09% 5.99%

Pipes and Valves 5.53% 6.03% 6.57% 6.80% 6.23%

Electrical 10.91% 5.98% 5.73% 5.86% 7.12%

Housing 8.63% 5.54% 5.41% 3.89% 5.87%

Contingencies 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component

Capital Cost

Category

Average 

Percent

International Consultants, Inc.

Contract 68-C6-0039 C-15

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C16.1 - Base Costs Obtained from the WATERCO$T Model for Lime Feed with Recalcination

Feed Capacity (lb/hr)

1,000 10,000

Manufactured Equipment $48,870 $80,660 p

Labor $1,510 $3,060 c

Pipes and Valves $3,120 $6,250 p

Electrical $6,880 $12,320 p

Housing $9,450 $26,250 p

Contingencies $10,470 $19,280 c

Total $80,300 $147,820

Table C16.2 - WATERCO$T Model Base Construction Cost Analysis for Lime Feed with Recalcination

Feed Capacity (lb/hr)

1,000 10,000

Manufactured Equipment 60.86% 54.57% 57.71%

Labor 1.88% 2.07% 1.98%

Pipes and Valves 3.89% 4.23% 4.06%

Electrical 8.57% 8.33% 8.45%

Housing 11.77% 17.76% 14.76%

Contingencies 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-16

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C17.1 - Base Costs Obtained from the WATERCO$T Model for Permanganate Feed Systems

Feed Capacity (lb/day)

1 10 100 500

Manufactured Equipment $2,340 $2,600 $3,380 $5,220 p

Labor $480 $480 $540 $770 c

Pipes and Valves $970 $970 $970 $970 p

Electrical $3,190 $3,190 $3,190 $3,190 p

Housing $1,260 $1,580 $1,950 $2,940 p

Contingencies $1,240 $1,320 $1,500 $1,960 c

Total $9,480 $10,140 $11,530 $15,050

Table C17.2 - WATERCO$T Model Base Construction Cost Analysis for Permanganate Feed Systems

Feed Capacity (lb/day)

1 10 100 500

Manufactured Equipment 24.68% 25.64% 29.31% 34.68% 28.58%

Labor 5.06% 4.73% 4.68% 5.12% 4.90%

Pipes and Valves 10.23% 9.57% 8.41% 6.45% 8.66%

Electrical 33.65% 31.46% 27.67% 21.20% 28.49%

Housing 13.29% 15.58% 16.91% 19.53% 16.33%

Contingencies 13.08% 13.02% 13.01% 13.02% 13.03%

Total 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-17

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C18.1 - Base Costs Obtained from the WATERCO$T Model for Polymer Feed Systems

Feed Capacity (lb/hr)

1 10 100 200

Manufactured Equipment $11,670 $11,670 $14,730 $18,970 p

Labor $700 $700 $700 $760 c

Pipes and Valves $280 $280 $280 $300 p

Electrical $1,290 $1,290 $1,290 $1,290 p

Housing $3,600 $3,600 $4,050 $4,500 p

Contingencies $2,630 $2,630 $3,160 $3,870 c

Total $20,170 $20,170 $24,210 $29,690

Table C18.2 - WATERCO$T Model Base Construction Cost Analysis for Polymer Feed Systems

Feed Capacity (lb/hr)

1 10 100 200

Manufactured Equipment 57.86% 57.86% 60.84% 63.89% 60.11%

Labor 3.47% 3.47% 2.89% 2.56% 3.10%

Pipes and Valves 1.39% 1.39% 1.16% 1.01% 1.24%

Electrical 6.40% 6.40% 5.33% 4.34% 5.62%

Housing 17.85% 17.85% 16.73% 15.16% 16.90%

Contingencies 13.04% 13.04% 13.05% 13.03% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-18

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C19.1 - Base Costs Obtained from the WATERCO$T Model for Rapid Mix, G=900

Basin Volume (ft3)

100 500 1,000 5,000 10,000 20,000

Excavation & Sitework $220 $380 $490 $1,360 $2,720 $5,460 c

Manufactured Equipment $4,310 $9,830 $14,760 $66,840 $133,670 $267,340 p

Concrete $390 $870 $1,280 $3,610 $7,220 $14,450 p

Steel $570 $1,350 $2,010 $5,600 $11,180 $22,360 p

Labor $1,230 $2,300 $3,410 $13,140 $26,280 $52,550 c

Electrical $6,980 $6,980 $7,180 $7,470 $8,760 $16,100 p

Contingencies $2,060 $3,260 $4,370 $14,700 $28,470 $56,740 c

Total $15,760 $24,970 $33,500 $112,720 $218,300 $435,000

Table C19.2 - WATERCO$T Model Base Construction Cost Analysis for Rapid Mix, G=900

Basin Volume (ft3)

100 500 1,000 5,000 10,000 20,000

Excavation & Sitework 1.40% 1.52% 1.46% 1.21% 1.25% 1.26% 1.35%

Manufactured Equipment 27.35% 39.37% 44.06% 59.30% 61.23% 61.46% 48.79%

Concrete 2.47% 3.48% 3.82% 3.20% 3.31% 3.32% 3.27%

Steel 3.62% 5.41% 6.00% 4.97% 5.12% 5.14% 5.04%

Labor 7.80% 9.21% 10.18% 11.66% 12.04% 12.08% 10.50%

Electrical 44.29% 27.95% 21.43% 6.63% 4.01% 3.70% 18.00%

Contingencies 13.07% 13.06% 13.04% 13.04% 13.04% 13.04% 13.05%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-19

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C20.1 - Base Costs Obtained from the WATERCO$T Model for Recarbonation, Liquid Carbon Dioxide

Installed Capacity (lb/day)

380 750 1,500 3,750 7,500 15,000

Manufactured Equipment $27,000 $31,000 $35,250 $49,250 $73,000 $141,000 p

Labor $7,650 $8,780 $12,170 $17,330 $28,990 $58,010 c

Pipes and Valves $1,530 $2,340 $4,620 $8,710 $16,940 $37,540 p

Housing $7,360 $7,360 $7,360 $7,360 $8,450 $8,900 p

Contingencies $6,530 $7,420 $8,910 $12,400 $19,110 $36,820 c

Total $50,070 $56,900 $68,310 $95,050 $146,490 $282,270

Table C20.2 - WATERCO$T Model Base Construction Cost Analysis for Recarbonation, Liquid Carbon Dioxide

Installed Capacity (lb/day)

380 750 1,500 3,750 7,500 15,000

Manufactured Equipment 53.92% 54.48% 51.60% 51.81% 49.83% 49.95% 51.93%

Labor 15.28% 15.43% 17.82% 18.23% 19.79% 20.55% 17.85%

Pipes and Valves 3.06% 4.11% 6.76% 9.16% 11.56% 13.30% 7.99%

Housing 14.70% 12.93% 10.77% 7.74% 5.77% 3.15% 9.18%

Contingencies 13.04% 13.04% 13.04% 13.05% 13.05% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-20

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C21.1 - Base Costs Obtained from the WATERCO$T Model for Recarbonation Basins

Single Basin Volume (ft3)

770 1,375 2,750 5,630 8,800 17,600 35,200

Excavation & Sitework $520 $620 $980 $1,390 $1,790 $3,050 $5,570 c

Concrete $1,380 $1,860 $2,820 $4,050 $5,190 $8,570 $15,320 p

Steel $2,250 $3,010 $4,670 $6,560 $8,320 $13,960 $25,240 p

Labor $2,830 $3,800 $5,730 $8,090 $10,240 $16,740 $29,730 c

Pipes and Valves $90 $130 $250 $480 $680 $1,360 $3,360 p

Contingencies $1,060 $1,410 $2,170 $3,090 $3,930 $6,550 $11,880 c

Total $8,130 $10,830 $16,620 $23,660 $30,150 $50,230 $91,100

Table C21.2 - WATERCO$T Model Base Construction Cost Analysis for Recarbonation Basins

Single Basin Volume (ft3)

770 1,375 2,750 5,630 8,800 17,600 35,200

Excavation & Sitework 6.40% 5.72% 5.90% 5.87% 5.94% 6.07% 6.11% 6.00%

Concrete 16.97% 17.17% 16.97% 17.12% 17.21% 17.06% 16.82% 17.05%

Steel 27.68% 27.79% 28.10% 27.73% 27.60% 27.79% 27.71% 27.77%

Labor 34.81% 35.09% 34.48% 34.19% 33.96% 33.33% 32.63% 34.07%

Pipes and Valves 1.11% 1.20% 1.50% 2.03% 2.26% 2.71% 3.69% 2.07%

Contingencies 13.04% 13.02% 13.06% 13.06% 13.03% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component

Capital Cost

Category

Average 

Percent

International Consultants, Inc.

Contract 68-C6-0039 C-21

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C22.1 - Base Costs Obtained from the WATERCO$T Model for Rectangular Clarifiers

Area (A=ft2) and Length x Width (LW=ftxft)

A=240

LW=30x8

A=600

LW=60x10

A=1260

LW=90x14

A=2240

LW=140x16

A=3600

LW=200x18

A=4800

LW=240x20

Excavation & Sitework $1,060 $2,000 $3,060 $4,680 $6,670 $8,090 c

Manufactured Equipment $8,540 $12,080 $24,470 $32,020 $53,110 $63,440 p

Concrete $2,970 $5,490 $8,430 $12,820 $18,190 $22,070 p

Steel $6,400 $13,110 $19,440 $32,620 $51,250 $69,680 p

Labor $6,220 $11,260 $17,320 $26,390 $37,570 $45,300 c

Pipes and Valves $6,960 $7,400 $9,100 $12,500 $16,100 $21,450 p

Electrical $1,510 $1,760 $1,860 $2,020 $2,110 $2,400 p

Contingencies $5,050 $7,970 $12,550 $18,460 $27,750 $34,860 c

Total $38,710 $61,070 $96,230 $141,510 $212,750 $267,290

Table C22.2 - WATERCO$T Model Base Construction Cost Analysis for Rectangular Clarifiers

Area (A=ft2) and Length x Width (LW=ftxft)

A=240

LW=30x8

A=600

LW=60x10

A=1260

LW=90x14

A=2240

LW=140x16

A=3600

LW=200x18

A=4800

LW=240x20

Excavation & Sitework 2.74% 3.27% 3.18% 3.31% 3.14% 3.03% 3.11%

Manufactured Equipment 22.06% 19.78% 25.43% 22.63% 24.96% 23.73% 23.10%

Concrete 7.67% 8.99% 8.76% 9.06% 8.55% 8.26% 8.55%

Steel 16.53% 21.47% 20.20% 23.05% 24.09% 26.07% 21.90%

Labor 16.07% 18.44% 18.00% 18.65% 17.66% 16.95% 17.63%

Pipes and Valves 17.98% 12.12% 9.46% 8.83% 7.57% 8.02% 10.66%

Electrical 3.90% 2.88% 1.93% 1.43% 0.99% 0.90% 2.01%

Contingencies 13.05% 13.05% 13.04% 13.05% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component

Capital Cost

Category

Average 

Percent

International Consultants, Inc.

Contract 68-C6-0039 C-22

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C23.1 - Base Costs Obtained from the WATERCO$T Model for Reverse Osmosis

Plant Capacity (mgd)

1.0 10 100 200

Manufactured Equipment $474,210 $3,458,480 $29,174,260 $56,438,930 p

Labor $70,420 $346,850 $2,312,340 $2,837,870 c

Electrical $65,740 $486,270 $3,635,690 $6,947,480 p

Housing $64,260 $462,650 $2,409,660 $4,176,740 p

Contingencies $101,190 $713,140 $5,629,790 $10,560,150 c

Total $775,820 $5,467,390 $43,161,740 $80,961,170

Table C23.2 - WATERCO$T Model Base Construction Cost Analysis for Reverse Osmosis

Plant Capacity (mgd)

1.0 10 100 200

Manufactured Equipment 61.12% 63.26% 67.59% 69.71% 65.42%

Labor 9.08% 6.34% 5.36% 3.51% 6.07%

Electrical 8.47% 8.89% 8.42% 8.58% 8.59%

Housing 8.28% 8.46% 5.58% 5.16% 6.87%

Contingencies 13.04% 13.04% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-23

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C24.1 - Base Costs Obtained from the WATERCO$T Model for Sodium Hydroxide Feed Systems

Feed Capacity (lb/day)

10 100 1,000 10000

Manufactured Equipment $6,440 $7,010 $5,720 $19,450 p

Labor $640 $640 $790 $4,120 c

Pipes and Valves $850 $850 $850 $850 p

Electrical $3,190 $3,190 $3,190 $3,460 p

Housing $1,010 $2,100 $8,400 $48,380 p

Contingencies $1,820 $2,070 $2,840 $11,440 c

Total $13,950 $15,860 $21,790 $87,700

Table C24.2 - WATERCO$T Model Base Construction Cost Analysis for Sodium Hydroxide Feed Systems

Feed Capacity (lb/day)

10 100 1,000 10,000

Manufactured Equipment 46.16% 44.20% 26.25% 22.18% 34.70%

Labor 4.59% 4.04% 3.63% 4.70% 4.24%

Pipes and Valves 6.09% 5.36% 3.90% 0.97% 4.08%

Electrical 22.87% 20.11% 14.64% 3.95% 15.39%

Housing 7.24% 13.24% 38.55% 55.17% 28.55%

Contingencies 13.05% 13.05% 13.03% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-24

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C25.1 - Base Costs Obtained from the WATERCO$T Model for Sulfuric Acid Feed Systems

Feed Capacity (gpd)

10 100 1000 5000

Manufactured Equipment $1,560 $3,440 $12,400 $41,000 p

Labor $640 $820 $2,840 $11,840 c

Pipes and Valves $1,090 $1,090 $2,150 $2,150 p

Electrical $1,670 $2,920 $2,920 $2,920 p

Housing $2,520 $1,560 $1,560 $1,560 p

Contingencies $1,120 $1,470 $3,280 $8,920 c

Total $8,600 $11,300 $25,150 $68,390

Table C25.2 - WATERCO$T Model Base Construction Cost Analysis for Sulfuric Acid Feed Systems

Feed Capacity (gpd)

10 100 1000 5000

Manufactured Equipment 18.14% 30.44% 49.30% 59.95% 39.46%

Labor 7.44% 7.26% 11.29% 17.31% 10.83%

Pipes and Valves 12.67% 9.65% 8.55% 3.14% 8.50%

Electrical 19.42% 25.84% 11.61% 4.27% 15.28%

Housing 29.30% 13.81% 6.20% 2.28% 12.90%

Contingencies 13.02% 13.01% 13.04% 13.04% 13.03%

Total 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-25

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C26.1 - Base Costs Obtained from the WATERCO$T Model for Tube Settling Modules

Tube Module Area (ft2)

280 2,800 14,000 28,000 56,000

Manufactured Equipment $4,200 $31,000 $147,000 $282,000 $504,000 p

Steel $2,000 $19,500 $95,000 $155,000 $300,000 p

Labor $2,500 $11,200 $49,000 $95,000 $224,000 c 

Contingencies $1,300 $9,300 $43,700 $79,800 $154,200 c

Total $10,000 $71,000 $334,700 $611,800 $1,182,200

Table C26.2 - WATERCO$T Model Base Construction Cost Analysis for Tube Settling Modules

Tube Module Area (ft2)

280 2,800 14,000 28,000 56,000

Manufactured Equipment 42.00% 43.66% 43.92% 46.09% 42.63% 43.66%

Steel 20.00% 27.46% 28.38% 25.34% 25.38% 25.31%

Labor 25.00% 15.77% 14.64% 15.53% 18.95% 17.98%

Contingencies 13.00% 13.10% 13.06% 13.04% 13.04% 13.05%

Total 100.00% 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-26

February 1999 

W/W Cost Model Capital Cost Breakdown 
Table C27.1 - Base Costs Obtained from the WATERCO$T Model for Wash Water Surge Basins

Capacity (gal)

10,000 50,000 100,000 500,000

Excavation & Sitework $200 $520 $1,250 $4,400 c

Concrete $11,560 $39,310 $71,480 $143,680 p

Steel $7,990 $25,170 $44,680 $70,770 p

Labor $18,270 $58,500 $107,590 $182,150 c

Pipes and Valves $5,500 $7,500 $11,000 $16,000 p

Electrical $1,300 $1,300 $6,000 $6,000 p

Contingencies $6,720 $19,850 $36,300 $63,450 c

Total $51,540 $152,150 $278,300 $486,450

Table C27.2 - WATERCO$T Model Base Construction Cost Analysis for Wash Water Surge Basins

Capacity (gal)

10,000 50,000 100,000 500,000

Excavation & Sitework 0.39% 0.34% 0.45% 0.90% 0.52%

Concrete 22.43% 25.84% 25.68% 29.54% 25.87%

Steel 15.50% 16.54% 16.05% 14.55% 15.66%

Labor 35.45% 38.45% 38.66% 37.44% 37.50%

Pipes and Valves 10.67% 4.93% 3.95% 3.29% 5.71%

Electrical 2.52% 0.85% 2.16% 1.23% 1.69%

Contingencies 13.04% 13.05% 13.04% 13.04% 13.04%

Total 100.00% 100.00% 100.00% 100.00% 100.00%

Cost Component

Cost Component Average 

Percent

Capital Cost

Category

International Consultants, Inc.

Contract 68-C6-0039 C-27

February 1999 

W/W Cost Model Capital Cost Breakdown 
This page was intentionally left blank.

International Consultants, Inc.

Contract 68-C6-0039 C-28

February 1999 

W/W Cost Model Capital Cost Breakdown 
Appendix D Basis for Revised Activated

Alumina Costs

D-1

APPENDIX D: BASIS FOR REVISED ACTIVATED ALUMINA COSTS

D.1 OVERVIEW

Four sets of costs were developed for disposable activated alumina (no regeneration) - two

for unadjusted pH and two where the pH has been adjusted to the optimal pH of 6. The costs for pH

adjustment to optimal pH 6.0 are included as a separate module in these cost estimates since it is only

used for two of the options. The costs in the cost models routinely used to estimate unit costs are all

based on regeneration of the media, operation at optimal pH, and use of single columns. Most of the

costs are based on fluoride removal rather than arsenic removal. It was determined that the existing

cost models could not be used to estimate costs for disposable activated alumina. A similar summary

to this was prepared for the proposed rule for the disposable activated alumina costs. It was included

in Appendix G of the November 1999 Technology and Cost Document (1). The main change between

that design and the design used for the final rule is that smaller columns operated in series will be used

instead of a single column. This will provide greater utilization of the media before disposal and is

more consistent with the designs used by commenters in evaluating disposable activated alumina.

D.2 DESIGN ASSUMPTIONS AND BASIS

D.2.1 Capital Costs - Natural pH

 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 ranges will be

evaluated 7.0 #pH < pH 8.0 and 8.0 # pH #8.3.

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
D-2

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. Since the run length is sensitive to pH, two ranges were developed for the

natural pH option.

 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(2). 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 (3).

The average raw water pH values for these two systems are pH 8.1 and pH 8.3 (3) Samples

were taken at the inlet and outlet to the activated alumina plants. Average pH and alkalinity

values at the inlet and outlet to the two plants were approximately equal. Since it does not

appear that activated alumina is adversely affecting the finished water pH, additional posttreatment corrosion control is not necessary.

 3. Empty Bed Contact Time (EBCT) = 5 minutes per column

Basis. Series operation with one column serving as a roughing column and the second column

serving as the polishing column provides greater utilization of the activated alumina than a

single column. Since the media is being operated on a disposable basis rather than being

regenerated, greater utilization of the media is desired to reduce operation and maintenance

(O&M) costs and minimize waste disposal. Since the volume of water treated is dependent

upon the size of the bed, an EBCT of 5 minutes was selected based on a comparison with other

designs. An EBCT of approximately 5 minutes was used for series operation with

regeneration in the AwwaRF report (4). The Phoenix case study was the only one to use

disposable media and the EBCT was 2.5 minutes (3). The EBCT in the White Rock Water
D-3

Company case study is 18 minutes (2). One full-scale site is based on a 7.5-minute EBCT and

the other is based on a 4.3 minute-EBCT (3).

 4. The estimated cost of the activated alumina media is $0.82/lb for systems with a design flow

greater than 0.1 mgd. The estimated cost of activated alumina media is $1.30/lb for systems

with a design flow less than or equal to 0.1 mgd.

Basis. The five case studies with regeneration in the AwwaRF report use $0.50/lb as the cost

of activated alumina media (3). The Phoenix case study uses $0.75/lb(4). Several

manufacturers and distributors were contacted for quotes on activated alumina. A range of

costs was obtained based on the size of the shipment. The lowest cost was based on a

truckload of 20 tons. The truckload costs averaged $0.68/lb. A higher cost of $0.82/lb was

assumed to account for higher costs that may be associated with supersack purchases at the

lower end of the flow range. The highest cost of $1.30/lb was for a 55 gallon barrel of 28x48

media. One of the full-scale plants uses Alcoa CPN 28x48 (fine) alumina and it cost $1.98/lb

when it was replaced in May 1999 (3). The other full-scale plant uses Alcoa DD-2 14x28

(coarse) alumina and it cost $1.01/lb when it was replaced in December 1998 (3). The barrel

cost was assumed for all very small systems with design flows at or below 0.1 mgd.

 5. The density of the activated alumina media is assumed to be 47 lb/ft3.

Basis. The bulk density of Alcoa CPN activated alumina is 47 /b/ft3 (3). The bulk density

range for Alcoa DD-2 activated alumina is 39 to 52 lb/ft3

 (3). The average for the DD-2

media would be 45.5 lb/ft3. The higher density was conservatively used to develop the costs.

 6. The bed depth assumption ranged 3 feet to 6 feet depending upon the design flow. The

maximum diameter per column was 12 feet. When the calculated media volume exceeded the

maximum, the treatment media was divided into several beds to maintain a realistic bed

diameter. Thus, larger designs include multiple activated alumina trains, each treating a

portion of the total flow.
D-4

Basis. An old activated alumina design manual for fluoride removal recommends that the bed

diameter be equal to or greater than the bed depth to prevent “wall effects” (5). This source

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 (6). Many of the conceptual designs in the cost estimating manual had bed depths that

exceeded the diameter, so the recommendation in the design manual was not strictly applied

to all designs. For design flows less than 50,000 gallons per day, the bed depth is greater than

the bed diameter. The data from the three case studies where activated alumina is being

operated without regeneration, the bed depth is greater than the bed diameter (2, 3).

 7. The vessel cost has been sized based on 50% bed expansion during backwash even though

backwashing may not be necessary on a routine basis for smaller systems. The vessel volume

was calculated as (1.5)(media volume).

Basis. The old design manual and the Water Model used 50% bed expansion to size the vessel

(5, 6). Both of these designs are based on regeneration of the media rather than disposable

operation. The bed expansion for the one case study where it could be calculated was 30%

(3). The design approach used as the basis for the case studies in the AwwaRF report used

80% expansion (7). However, this design was also based on regeneration of the media rather

than disposable media operation and was based on designs greater than 1 mgd. Since the

expansion is only necessary for backwashing, 50% expansion was assumed for all sizes.

Backwashing is not performed on a routine basis at any of the three small systems operating

activated alumina without regeneration.

 8. The vessel cost is based on the following equation: 

Cost = 63.288 * (Vessel Volume in gallons) ^ 0.679

Basis. Quotes were obtained from two manufacturers for glass-lined carbon steel vessels with

working pressure rating between 75 and 150 psi and meeting American Society of Mechanical

Engineers (ASME) Code Section VIII requirements. Sizes ranging from 140 gallons to 6764

gallons were used to develop the equation. Vessel costs for sixteen different sizes were
D-5

regressed against their respective volumes to derive the equation. Costs from the AwwaRF

report cannot be used for comparison because of an apparent error in the vessel costs. The

vendor quote used for vessel costs is actually a vendor quote for a complete activated alumina

system. The vendor quote presents budget prices for three design flows and different size

vessels are used for each design flow. The vessel sizes are listed with the budget price and

this is apparently the source of the mistake. The three full-scale systems using activated

alumina on a disposable basis use fiberglass tanks instead of carbon steel tanks because the

pressure drop is very small. Fiberglass tanks would be less expensive than the carbon steel

vessels, so the vessel costs are conservative for very small systems.

 9. The capital costs include a redundant column to allow the system to operate while the media

is being replaced in the old roughing column.

Basis. A redundant vessel has been included for each set of roughing and polishing columns.

While the activated alumina media is being replaced in this column, the other two columns can

still function as a roughing and polishing column. Activated alumina media is not included in

the redundant vessel(s) as a capital cost to make it easier to estimate O&M costs. The Phoenix

case study uses a redundant roughing and polishing column (4). This design did not seem to

be the most efficient because one column would contain partially used activated alumina, but

would not be in operation while the media in the other column was being replaced. The initial

Albuquerque design did not contain a redundant column, but instead relied upon the old

polishing column to treat all the water while the media was replaced in the old roughing

column (7). The two full-scale sites in New Hampshire split the flow into two streams and

had roughing and polishing columns treating half the water (3). Redundant columns were not

included; however, one set of columns could treat all the water if there was a problem with

the other set.

 10. The remaining components of the process equipment include concrete, pipe and valves and

electrical and instrumentation. A cost factor of 14.43% of the manufactured equipment and

activated alumina was used for these other construction cost components.
D-6

Basis. The Water Model provided construction cost estimates for a variety of conceptual

designs (6). Construction cost components included process costs such as manufactured

equipment and media, concrete, pipes and valves, electrical and instrumentation as well as

construction costs such as excavation and sitework. The conceptual designs are based on an

EBCT of 7.5 minutes. Since this design is using a total EBCT of 10 minutes (two 5-min

EBCTs in series), the activated alumina costs were adjusted 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 10

minutes. Concrete was 3.17%, pipe and valves were 15.18%, and electrical and

instrumentation were 11.12% 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 2. The electrical

and instrumentation were reduced by a factor of 3 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 14.43%. 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. For the designs

with pH adjustment, a separate module has been developed including pipes and valves for

those chemical feed systems.

 11. The capital costs have been estimated from the total process costs using a factor of 2.5 for

small systems and 3.33 for large systems.

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 (8). The process costs

were assumed to be 40% of the capital costs for small systems. The breakdown of capital
D-7

costs for small systems is as follows: 40% process costs, 40% construction costs, and 20%

engineering costs. For large systems, the breakdown of capital cost is: 30% process costs,

40% construction, and 30% engineering.

 12. 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. 

Basis. Data from the 1995 Community Water Systems Survey were used to identify the

percentage of systems without treatment (9). The data from ground water systems is as

follows:

Table D-1

Percentage of Ground Water Systems with No Treatment: Systems Serving # 3,300 People

Population Category

25 - 100 101 - 500 501 - 1,000 1,001 - 3,300

43% 19% 16% 18%

Table D-2

Percentage of Ground Water Systems with No Treatment: Systems Serving > 3,300 People

Population Category

3,301 - 10,000 10,001 - 50,000 50,001 - 100,000 100,001 - 1 MIL

13% 1% 11% 0%

Larger ground water systems rely on multiple entry points to supply ground water rather than

one well. Thus, treatment may be present in the system, but only at one of the wells. Since the

analysis is based on no treatment throughout the system, using the percentages for larger
D-8

systems could underestimate the need for roads and fences at some entry points. Therefore,

the 18% without treatment from the 1,001 - 3,300 category was applied to the four larger

system categories. The following table summarizes the percentage used for each of the size

categories.

Table D-3

Assumed Percentage of Ground Water Systems with No Treatment

Population Category Percentage of GWS w/o any Treatment

25 - 100 43%

101 - 500 19%

501 - 1,000 16%

1,001 - 3,300 18%

3,301 - 10,000 18%

10,001 - 50,000 18%

50,001 - 100,000 18%

100,001 - 1 MIL 18%

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.

 13. 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 = 726.6 * Design Flow + 59

Basis. The basis for the cost equations for buildings, fence and road are described in Section

4.0 of the Very Small Systems Document (10). The Very Small Systems Document provided

costs as a percentage of the capital costs for design flows from 0.024 to 0.65 mgd. The

process area was calculated from the building cost estimates by dividing by 40 (building costs
D-9

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:

Process Area (sq ft) = 614.6 * Design Flow + 68

The process area for the conceptual designs in the Water Model (6) were also regressed and

fitted to a linear regression. The conceptual designs covered a design flow range from 0.045

to 1.082 mgd. The linear regression from this source is:

Process Area (sq ft) = 726.6 * Design Flow + 59

The process area estimates using the regression of the older model data were selected to make

cost estimates for buildings, road and fence. Since there would be a minimum building size,

extrapolating a linear regression below the lowest data point does not create a significant

problem with the intercept. These process area estimates were larger than the process area

estimates from the Very Small Systems Document for design flows larger than 0.09 mgd. At

the lower design flows, the process area estimates were slightly lower using the Water Model.

The Water Model covered a broader range of design flows, so less extrapolation is required

at higher design flows than would be the case if the Very Small Systems Document equation

were used. 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 compared with the minimum area needed for the

vessels using the vessel diameter. The process area estimates are greater than three times the

minimum area estimates for the vessels, including the redundant column. 

 14. The building cost to house the activated alumina system is $45.38/sq ft for design flows at or

below 0.65 mgd. The following equation is used for design flows between 0.65 and 1 mgd:

Building cost = 54.204 * Process Area + 14887
D-10

For design flows greater than 1 mgd, the building cost varied based on flow from $100 to $90

per square foot.

Basis. The data from both the Very Small Systems Document and 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 (10). 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 prefabricated 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 (6). 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 ENR Building Cost Index was used to update both sets of costs to

September 1998. The housing rate from the Very Small Systems Document was selected

based on better documentation and the use of more recent data for design flows up to 0.65

mgd. For design flows between 0.65 and 1 mgd, a linear regression of housing cost against

process area was used to derive the equation listed above. For design flows greater than1

mgd, the building cost rate assumptions were the same as those used in the for Albuquerque

Report (7) and for the case studies in the AwwaRF Report (4). 

D.2.1 Capital Costs - pH Adjustment

 15. For two options, pH adjustment to the optimal pH 6 is utilized. The capital cost equation for

the pH adjustment module is:

Capital cost = 40517 * Design Flow + 31889
D-11

Basis. Capital cost components include piping, acid and caustic metering pumps, and storage

tanks for the acid and caustic. For the metering pumps, pH controllers, pH sensors, and level

sensors are included in the costs. Cost data for each of the components were obtained from

manufacturers for design flows ranging from 0.1 to 3 mgd. Capital costs were estimated by

multiplying the total process cost by 2.5 or 3.33 depending upon size(8). The Phoenix case

study in the AwwaRF report provides the only realistic comparison for disposable activated

alumina at optimal pH (4). The other case studies are based regeneration and would use acid

and caustic for those purposes. A design flow of 0.2 mgd was evaluated in both the Phoenix

case study and the process cost estimate used to derive the above equation. The process cost

estimate of $16,200 is very similar to the process cost from the Phoenix case study of $16,750.

D.2.1 Operation and Maintenance Costs 

There are three major components to the operation and maintenance (O&M) costs -

replacement media, energy and labor.

 16. 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. The run

length for the natural pH range of 7 # pH < 8 is 10,000 bed volumes (BV). The run length for

the natural pH range of 8 # pH # 8.3 is 5,200 BV. These run lengths are used to estimate the

replacement of the activated alumina in half of the active columns. 

Basis. The full-scale systems in New Hampshire serve as the basis for the run lengths for the

natural pH options (3). Plant C data were used for the range of 7 # pH < 8 and data from Plant

D were used for the range of 8 # pH # 8.3. 
D-12

The source water at Plant C has the following critical characteristics:

Parameter Average Source Water Results

Arsenic (total) 59 Fg/L

pH 8.1

Alkalinity 84 mg/L CaCO3

Fluoride 1.5 mg/L

Sulfate 26 mg/L

The roughing tank treated 10,050 BV of water before breakthrough (effluent = influent). The

arsenic concentration in the effluent from the second tank was approximately 20 Fg/L. Based

on the sample results from September 2 and September 16, 1998, breakthrough above 3 Fg/L

(95% removal based on average inlet of 62.1 Fg/L) in the effluent occurred somewhere within

that time period. This converts to a run length of approximately 8500 BV. The roughing tanks

were replaced with virgin activated alumina after treating 11,071 BV of water. The former

polishing tanks then became the roughing tanks. The data indicate that the outlet arsenic

concentration was still below 4 Fg/L as the total run approached 16,000 BV. This supports

a run length of 8500 BV per replacement of half of the media. Arsenic concentration was the

only key influent parameter that varied to any extent during the testing.

Run length would be affected by pH and by the influent arsenic concentration. Since influent

arsenic concentrations for most systems affected by the rule should be much lower than 59

Fg/L and since arsenic removal is pH dependent, a run length of 10,000 BV was selected for

the range of 7 # pH < 8. This should be very conservative for systems with lower arsenic

concentrations and pH values at the lower end of the range. The effluent pH and effluent

concentrations of alkalinity, fluoride and sulfate indicated almost no change from their value

in the influent. Thus, the existing corrosion control should be sufficient.
D-13

The source water at Plant D has the following critical characteristics:

Parameter Average Source Water Results

Arsenic (total) 59.0 Fg/L

pH 8.3

Alkalinity 55.5 mg/L CaCO3

Fluoride 1.13 mg/L

Sulfate 15 mg/L

The roughing tank treated 5,200 BV of water before breakthrough (effluent = influent). The

arsenic concentration in the effluent from the second tank was less than 2 Fg/L as it was

throughout the run. Arsenic concentration was the only key parameter that varied to any extent

during the testing. The run length would be affected by pH and by the influent arsenic

concentration. The pH was consistent and represents the upper bound on the range of

applicability for this cost equation. A run length of 5,200 BV was used as the run length for

the range of 8 # pH # 8.3. Since influent arsenic concentrations for most systems affected by

the rule should be much lower than 59 Fg/L, this should be a very conservative run length for

systems with lower arsenic concentrations and/or pH values at the lower end of the range.

The effluent pH and effluent concentrations of alkalinity, fluoride and sulfate indicated almost

no change from their value in the influent. Thus, the existing corrosion control should be

sufficient.

The previous approach described in Appendix G of the November 1999 Technology and Cost

Document (1) relied upon pilot-scale data from Severn-Trent (11). The study examined

arsenic removal over a pH range from pH 6 to pH 7.5 in a single column. The three run

lengths derived from that data ranged from 16,500 BV and pH 7 to 3,000 BV at pH 8 for 10%

breakthrough in a single column. Actual run lengths to 50% breakthrough in that study for

natural pH 7.5 ranged from 9,400 to 13,000 BV based on mesh size. The data from the two

full-scale facilities (3) indicated that the roughing column could be operated almost complete

breakthrough (effluent = influent) before concentrations in the effluent of the second column

exceeded 5% of the influent concentration. This data also indicates that a natural pH run
D-14

length of 10,000 BV for 7 # pH < 8 may be conservative, especially for systems with lower

pH.

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.

 17. The two previous runs were based on operating at the natural pH without pH adjustment.

Adjusting pH in larger systems is less likely to be an issue, so run lengths were also

developed for operation at pH 6. Two run lengths were selected for the optimal pH of 6 -

15,400 and 23,100 BV. 

Basis. The two run lengths for pH adjustment to pH 6 are based on pilot-scale data performed

at Albuquerque (12). The Arsenic Treatment Evaluation Report (7) used the pilot-scale data

to derive a run length estimate for series operation of activated alumina columns with

regeneration. In the examination of this report, it was determined that there was an error in

the derivation of the run length estimate (13). A revised estimate of run length was calculated

using the same data. The procedure described in that following modified excerpt describes

the procedure used to calculate the run length at pH 6 for disposable activated alumina. 

 A. The equation for arsenic effluent in the Arsenic Treatment Evaluation Report (7) was used to

determine the adsorptive capacity of the lead column until arsenic breakthrough at 2 Fg/L. The

equation was based on the data from Figure 14 of the Phase 3 Report (12). It was necessary

to make one assumption to estimate the arsenic adsorbed onto the column. It was assumed that

the influent concentration of 22.8 Fg/L was constant. From the Phase 3 Report, one column

can treat 4300 BV before arsenic breaks through at 2 Fg/L. The amount of arsenic that passed

through the column can be determined by estimating the area under the curve defined by the
D-15

equation y = (2*10-8)x2 + 0.0002x + 0.8394. The integral of this equation was used to estimate

the area under the curve. The integral equation is:

z = (2*10-8)/3*x3 + 0.0001x2 + 0.8394x. 

For x = 4300 BV:

Aspass = (2*10-8)/3*(4300)3 + 0.0001*(4300)2 + 0.8394*4300

Aspass = 5988 Fg*BV/L.

The arsenic adsorbed onto the activated alumina was calculated by subtracting the arsenic that

passed through the column from the total arsenic in the influent. 

Asads = Astot - Aspass = (4300 BV)(22.8 Fg/L) - 5988 Fg*BV/L

Asads = 98,040 - 5988 Fg*BV/L = 92,052 Fg*BV/L

The arsenic adsorbed onto the activated alumina is often reported in milligrams arsenic per

gram of activated alumina. The Phase 3 Report refers to this as As loading in Table 4. In the

Phase 3 Report, a bed volume was 0.375 L. To calculate the arsenic loading, the density of

Alcoa CPN is needed. The packed bulk density for Alcoa CPN is 47 lb/ft3 or 750 kg/m3.

Arsenic loading is calculated as follows:

As loading = [(92,052 Fg*BV/L)(0.375 L/BV)(mg/1000 Fg)]

 (0.375 L)(dm3/L)(m3/1000 dm3)(750 kg/m3)(1000 g/kg)

As loading = 0.12 mg As/g AA for breakthrough at 2 Fg/L

Arsenic loading will vary based on the breakthrough concentration. For breakthrough at 10

Fg/L, the run length is 16,500 BV. Using the integral equation yields the following:

Aspass = 71,023 Fg*BV/L

Asads = (22.8 Fg/L)(16,500 BV) - 71,023 Fg*BV/L = 305,177 Fg*BV/L
D-16

The arsenic loaded onto the column is:

As loading = [(305,177)(0.375)(1/1000)] 

(0.375)(1/1000)(750)(1000)

As loading = 0.41 mg As/g AA {Table 2 of Phase 3 Report lists 0.43 mg As/g AA}

 B. Using the bed volume size of 16,032 liters, yields the following adsorptive capacities:

Asads = (92,052 Fg*BV/L)(16,032 L/BV)(1 gram/106 Fg) = 1476 grams

Astot = (98,040)(16,032)(1/106) = 1572 grams

It appears that the adsorptive capacity of 1573 grams in the Arsenic Treatment Evaluation

Report is based on total arsenic instead of the adsorbed arsenic. 

 C. The maximum run length will be based on an adsorptive capacity of 1476 grams in the second

column. For this step of the analysis, it will be assumed that all of the arsenic passing through

the first column will be adsorbed by the second column. This is a very conservative estimate

because the second column will also have leakage through it and concentrations will rise until

breakthrough is reached at 2 Fg/L. This is balanced by the fact that when the polishing column

is moved into the roughing column position, some of its capacity will have already been used

when it was the polishing column. Both of these factors will be addressed later. The integral

equation in step 2 is again utilized to estimate the amount of arsenic passing through the first

column. When this value is equal to 1476 grams, then the maximum run length can be

determined.

For a run length of 18,500 BV:

Aspass = (2*10-8)/3*(18500)3 + 0.0001*(18500)2 + 0.8394*18500

Aspass = 1474 grams
D-17

The maximum run length is 18,500 bed volumes. The amount of arsenic adsorbed onto the first

column is:

Asads = (22.8 Fg/L)(18,500 BV)(16,032 L/BV)(g/106 Fg) - 1474 grams

Asads = 5288 grams

Thus, the capacity of the roughing column before it needs to be regenerated is 5288 grams.

This assumes that there is no arsenic already adsorbed onto the column. Under steady-state

operation, the polishing column will become the roughing column and the redundant column

with virgin media column becomes the polishing column. The media in the old roughing

column is replaced with virgin media and becomes the new redundant column. The adsorbed

arsenic on the former polishing column will decrease run length whereas arsenic leakage

through the second column will increase run length.

 D. The adsorptive capacity of the roughing column was 5288 grams based on the maximum run

length of 18,500 bed volumes. When the “old” polishing column is shifted and becomes the

“new” roughing column, it already has a maximum of 1474 grams adsorbed onto it. Therefore,

the remaining capacity is considerably lower than 5288 grams. The absorptive capacity of the

“new” roughing column was first estimated by subtracting the arsenic that passed through the

first column. This estimate was called the minimum capacity of column 2, since the polishing

column does not adsorb all of the arsenic that passed through the first column.

Min Adsorptive Capacity - Column 2 = 5288 - 1474 grams = 3814 grams.

 E. As noted in the previous step, the minimum adsorptive capacity was estimated by assuming

that all of the arsenic that passed through the first column was adsorbed onto the second

column. This would not be the case. The second column would have leakage effects as well

as increasing arsenic concentrations until the effluent concentration reached 2 Fg/L. Because

the arsenic passing through the first column and reaching the second column is not constant,

it is not possible to estimate a breakthrough curve. Therefore, a different approach was taken

to better estimate the adsorptive capacity. Arsenic leakage from the single column study

summarized in the Phase 3 Report (12) was examined to estimate leakage for the second
D-18

column. The Phase 3 Report noted that arsenic leakage was always greater from the onceregenerated alumina than the virgin alumina. For the 5-minute EBCT runs, an average arsenic

leakage (up to 3200 BV) of 0.3 and 1.4 Fg/L was observed for the virgin and once-regenerated

Alcoa CPN alumina runs. For the Alcan alumina at an EBCT of 5 minutes, average arsenic

leakage (up to 3200 BV) was 0.2 and 1.3 Fg/L for the virgin and once-regenerated alumina

runs. After 3200 BV, the normal breakthrough curve started, so the effluent arsenic was not

solely due to leakage. Since the run length of 4300 BV and breakthrough equation were both

based on the once-regenerated Alcoa CPN alumina, arsenic leakage was also estimated for

this alumina. The average arsenic leakage was 1.4 Fg/L It was assumed that the second

column would exhibit the same leakage characteristics as the first column. Thus, arsenic

leakage for the second column was assumed constant at 1.4 Fg/L. It was assumed that 1.4

Fg/L would pass through both columns starting at the beginning of the run. The second column

will probably remove some of the arsenic that is passing through the first column, so this

approach likely overestimates the amount of arsenic leakage for the second column. However,

since there was no way to estimate a breakthrough curve for the second column, it was not

possible to estimate arsenic in the effluent due to reduced adsorptive capacity. The arsenic

leakage estimate was used to account for all arsenic that would pass through the second

column - from both leakage and reduced adsorptive capacity. Thus, it likely underestimates

the amount of arsenic that passed through the second column into the finished water. The

arsenic leakage was estimated using the following equation:

Asleakage = (1.4 Fg/L)(RL BV)(16,032 L/BV) Where RL = run length in BV

For a run length of 18,500 BV:

Asleakage = 415 grams

 F. The available arsenic capacity of the second column is estimated by adding the arsenic

leakage to the minimum arsenic capacity of the second column.

Available Arsenic Capacity of Column 2 = 3814 grams + 415 grams = 4229 grams
D-19

Since the available arsenic capacity of the “old” polishing column is lower than the maximum

capacity, the steady-state run length must be lower than 18,500 bed volumes.

 G. The run length was determined by an iterative approach. The integral equation was used to

estimate the arsenic adsorbed onto the first column for the lower run lengths. The available

arsenic capacity of the second column was also estimated using the arsenic leakage determined

for the new run length. When the arsenic adsorbed on column 1 was equivalent to the

available arsenic capacity of column 2, the steady-state run length was determined. The

following will illustrate this process:

At 15,400 bed volumes:

Column 1 - arsenic adsorbed = 4651 grams

Column 1 - arsenic passed through = 978 grams

Column 2 - arsenic leakage = 346 grams

Column 2 - arsenic adsorbed = 632 grams

Column 2 - available arsenic capacity = 5288 - 632 = 4656 grams

The available arsenic capacity of column 2 is approximately equivalent to the amount of

arsenic that can be adsorbed onto it when it becomes the roughing column (column 1) based

on the maximum run length. This is the steady-state run length to meet a standard of 2 Fg/L at

the effluent to the second column. After 15,400 bed volumes, the media in the roughing column

is taken off-line and replaced. While this is occurring, the polishing column becomes the new

roughing column and the redundant column becomes the new polishing column.

 H. The Phase 3 Report (12) recommended that the run length be increased by up to 50% to

account for improved performance when hydrochloric acid is used instead of sulfuric acid to

adjust the pH. Based on the Phase 3 Report, low and high estimates of run length were made

based on the use of hydrochloric acid for pH adjustment. For the low option, a 20% increase

was assumed and a 50% increase was assumed for the high option. If hydrochloric acid is

used for pH adjustment, the run length would range from 18,480 bed volumes to 23,100 bed

volumes.
D-20

The lower bound run length at pH 6 was based on the calculated run length for Albuquerque

water when the pH is adjusted using sulfuric acid (15,400). The upper bound run length at pH

6 was based on the higher end of the range when hydrochloric acid is used instead of sulfuric

acid. There are several factors that could influence run length.

The first factor is that the run length for the single column is based on the data from onceregenerated activated alumina. It was not possible to determine the effect of regeneration

because the influent concentration varied significantly between the two runs. Some studies

have shown some degradation after regeneration. The use of virgin media produced a longer

run length even though the influent concentration was much higher (38.8 Fg/L). Effluent

concentrations from virgin media did not exceed 2 Fg/L (.95% removal) until 7000 BV had

been treated using either the Alcoa CPN and the Alcan media. This is nearly double the single

unit run length of 4300 BV using the once regenerated media that serves as the basis for the

two pH adjusted series operation options. Using the run lengths based on regenerated media

is likely to be very conservative for costing disposable activated alumina options.

The second factor is that the leakage through each of the columns will be reduced. The data

described above for virgin media was less than 0.4 Fg/L. This may also extend the run length

for series operation based on virgin media because less arsenic is loaded onto the second

column due to leakage.

The third factor is that the data from the two full-scale facilities (3) indicated that the roughing

column could be operated almost complete breakthrough (effluent = influent) before

concentrations in the effluent of the second column exceeded 5% of the influent concentration.

Complete breakthrough refers to the effluent of the first column (C1) equaling the influent to

the column (C0). The finished water would be the effluent of the second column (C2). The

ratio of C1/C0 was derived for the Alcoa CPN data in Figure 14 of the Phase 3 Report (12).

C1 .15 Fg/L at the end of the data collection (.17,000 BV)

C0 = 38.8 Fg/L

C1/C0 = 0.39
D-21

If the first column can be operated until near complete breakthrough while maintaining arsenic

concentrations at or below 2 Fg/L, then the run lengths of 15,400 BV and 23,100 BV are

probably underestimated due to the conservative assumptions used in their derivation.

 18. The other major chemical cost is the pH adjustment chemicals for the two designs where pH

is adjusted to the optimal pH. The following equation was used for chemical costs for pH

adjustment:

Chemical cost (pH adjustment) = 81361* Average Flow + 151.7

Basis. The pH is adjusted downward using hydrochloric acid. The chemical demand was

calculated based on adjusting pH from pH 8 down to pH 6 with an alkalinity of 100 mg/L as

CaCO3. The pH is adjusted back to pH 8 using sodium hydroxide. Chemical costs were

developed for average flows ranging from 0.05 to 1.5 mgd. The Phoenix case study presented

chemical costs of $5742 for a flow of 0.2 mgd (4). The pH was adjusted from the pH 7.5 -

7.8 range down to pH 6 for those costs. The alkalinity was not listed for Phoenix. For an

average flow of 0.15 mgd, the chemical costs were estimate to be $12,208. The chemical

costs may be conservative for systems with lower pH and lower alkalinity. Systems with

higher alkalinity may experience higher costs. It was not possible to compare with the other

AwwaRF case studies because acid and base are used for regeneration as well as pH

adjustment.

 19. The second major component in the O&M costs is incremental labor cost. Incremental labor

is the labor associated with the additional maintenance that comes 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 5% breakthrough) depending upon

the volume of media being replaced. Sixteen hours was assumed for design flows less than

1 mgd and thirty-two hours for larger design flows. Incremental labor for the pH adjustment

processes is also included in the two options operated at optimal pH. The incremental labor

is an additional two hours per week.
D-22

Basis. The labor requirements when operated at natural pH 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 system (2). 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 doubled for the design flows greater than 1 mgd because of the larger

volume of media to be replaced. When operated at the optimal pH, an additional two hours

per week is used for the pH adjustment processes. Technical oversight is reduced due to the

use of pH controllers, pH sensors, and level sensors in this design.]

 20. The labor rate for small systems was $28/hour. The loaded labor rate for large systems was

$52/hour.

Basis. A rate of $28/hour is recommended as a loaded labor rate for small systems at the EPA

Technology Design Workshop Summary included in Annex A of the Phase I Upgrade

Document (8). A loaded labor rate of $52/hour was selected for large systems based on

licensed operators. For this analysis, a large system was defined as a system with an average

flow greater than 0.36 mgd (corresponds to a design flow of 1 mgd). This corresponds to

populations lower than 3,300 people. The use of a large system rate may overestimate costs

in systems serving between 2,700 and 10,000 people.

The Design Workshop Summary (8) also noted that the Bureau of Labor Statistics (BLS)

publishes monthly reports that include fully burdened labor costs on a regional basis. Other

members of the Panel noted that some States have conducted operator surveys that include

labor rates. The Panel recommended that EPA review and compare both the BLS and State

material to derive representative rates. A preliminary analysis of such material indicates that

the labor rates for both large and small systems may be very conservative and that actual labor

rates may be much lower.]

 21. 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 ft/yr,
D-23

which includes lighting, ventilation, and heating. Process electrical energy was assumed to

be 440 kwh/yr. Costs were based on a rate of $0.08/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 EC) 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. Power costs were consistently higher

using this approach compared with the Phoenix data (4).

D.3 WASTE DISPOSAL COSTS

 22. The primary waste is the spent activated alumina. Rates for disposal at sanitary landfills are

based on tons of material. A wet density is assumed to be 10% higher than the dry density to

convert the volume of activated alumina into a mass for disposal.

Basis. The cost estimate for the White Rock Water Company was based on the mass of the

media and did not assume a higher wet density (2). The Phoenix case study was the only one

that investigated disposable activated alumina. The disposal costs are based on a 10%

increase in the mass of the media (4). The wet density of the media was assumed to be 10%

higher than the dry density.
D-24

 23. The disposal cost for the spent activated alumina media is $60/ton.

Basis. The disposal rate of $60/ton was used to estimate costs for the White Rock Water

Company (2). The two water system byproducts documents (14, 15) contain the same cost

equation for off-site nonhazardous landfill cost for various system sizes. The cost equation

for disposal in a nonhazardous landfill from these two documents was also examined to

estimate disposal costs. That cost equation had two variables: tons of sludge requiring

disposal and transportation distance. The transportation distance could range from 5 to 50

miles. Using the upper bound of 50 miles, the equation simplified to a multiplier of

$45.58/ton. The Phoenix case study used a disposal rate of $120/wet ton (4). The basis for

this rate is not discussed. The $60/ton rate was selected for the cost analysis because it was

based on a full-scale system that was disposing spent alumina. The volume of waste being

generated is not significant, so even using $120/wet ton would not appreciably affect the waste

disposal costs.

 24. 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 (10) and the Water Model (6) provided a

backwash rate of 8 - 9 gpm/sq ft and a backwash duration of 10 minutes. The backwash rate

and duration are based on a design where the activated alumina media is being regenerated.

Backwashing was not performed at the White Rock Water Company probably due to the

frequent media replacement. Backwashing is not performed on a routine basis at the other two

full-scale systems (3). At Plant C, backwashing was performed twice over a period from

August 1997 to September 1999. When backwashing was performed on August 5, 1998, the

backwash rate was 3 gpm and the duration was 5 minutes. At Plant D, backwashing was

originally scheduled every four months, but this was discontinued after it was found to be

unnecessary. Backwashing was not performed between May 1998 and September 1999.

Backwashing costs are not included in the Phoenix case study, which is the only case study to

evaluate disposable activated alumina (4). Routine backwashing is not essential to operation

of these plants and if performed, would produces a very small waste stream. Backwash water
D-25

disposal costs, if needed, are covered by the conservative assumptions that have been used

throughout the costing process.

D.3 REFERENCES

 1. US EPA. Technologies and Costs for Removal of Arsenic from Drinking Water. Prepared

by International Consultants, Inc. and Malcolm Pirnie, Inc. under contract 68-C6-0039 for EPA

OGWDW. November 1999.

 2. Stewart, H.T., and K. J. Kessler. “Evaluation of Arsenic Removal by Activated Alumina

Filtration at a Small Community Public Water Supply,” Journal New England Water Works

Association. September 1991.

 3. US EPA. Arsenic Removal from Drinking Water by Ion Exchange and Activated Alumina

Plants. Prepared by Battelle under contract 68-C7-0008 for EPA ORD. October 2000.

 

 4. American Water Works Association Research Foundation. Cost Implications of a Lower

Arsenic MCL. Prepared by Frey, M.M. et al for American Water Works Association Research

Foundation. May 5, 2000.

 5. US EPA. Design Manual: Removal of Fluoride from Drinking Water Supplies by Activated

Alumina. Prepared by Rubel and Hagler, Inc. under contract 68-03-2917 for EPA ORD.

August 1984.

 6. US EPA. Estimation of Small System Water Treatment Costs, Prepared by

Culp/Wesner/Culp under contract 68-03-3093 for EPA ORD. 1983.

 7. CH2M Hill. Arsenic Treatment Evaluation Report. Prepared for the City of Albuquerque

Public Works Department. April 1999.
D-26

 8. US EPA. Guide for Implementing Phase I Water Treatment Cost Upgrades. Prepared by

Science Applications International Corporation under contract 68-C6-0059 for EPA

OGWDW. September 1998.

 9. US EPA. Geometries and Characteristics of Public Water Systems. Prepared by Science

Applications International Corporation under contract 68-C6-0059 for EPA OGWDW. August

15, 1999.

 10. US EPA. Very Small Systems Best Available Technology Cost Document. Prepared by

Malcolm Pirnie, Inc. for EPA. September 1993.

 11. Simms, J. and F. Azizian. “Pilot-Plant Trials on the Removal of Arsenic from Potable Water

Using Activated Alumina,” AWWA Water Quality Technology Conference, Denver, Colorado:

American Water Works Association, November 1997.

 12. Clifford, D.A. G. Ghurye et al. Final Report: Phase 3 City of Albuquerque Arsenic Study

Field Studies on Arsenic Removal in Albuquerque, New Mexico using the University of

Houston/EPA Mobile Drinking Water Treatment Research Facility. Houston, TX:

University of Houston, August 1998.

 13. US EPA. “Analysis of Activated Alumina Costs for Albuquerque.” January 28, 2000.

 14. US EPA. Small Water System Byproducts Treatment and Disposal Cost Document.

Prepared by DPRA Incorporated for EPA OGWDW. April 1993.

 15. US EPA. Water System Byproducts Treatment and Disposal Cost Document. Prepared by

DPRA Incorporated for EPA OGWDW. April 1993.
Appendix E Basis for Revised Anion Exchange

Costs
.
E-1

APPENDIX E: BASIS FOR REVISED ANION EXCHANGE COSTS

E.1 OVERVIEW

Two sets of costs were developed for anion exchange - one where the sulfate concentration

is less than or equal to 20 mg/L and the other where sulfate is greater than 20 mg/L, but less than or

equal to 50 mg/L. The costs in both the April 1999 Technology and Cost Document (1) and the

November 1999 Technology and Cost Document (2) utilize models based on nitrate removal. The

Water Model (3) was one of the cost models used for small systems with design flows less than 1

mgd. Sulfate concentrations for this model were between 50 and 100 mg/L. The November 1999

Technology and Cost Document listed the design criteria for large systems (> 1 mgd) using the W/W

model. The sulfate concentration was 80 mg/L and the nitrate concentration was 100 mg/L. It was

not possible to modify the sulfate concentration in either of these two models. It was determined that

the existing cost models could not be used to estimate costs for anion exchange options with sulfate

at or below 50 mg/L. A similar summary to this was prepared for the disposable activated alumina

costs used in the proposed and final rules.

E.2 DESIGN ASSUMPTIONS AND BASIS

E.2.1 Capital Costs

 1. Empty Bed Contact Time (EBCT) = 2.5 minutes per column

Basis. The designs in the November 1999 Technology and Cost Document were based on an

EBCT of 2.5 minutes. Case study costs in the AwwaRF report (4) were based on EBCTs of

1.5 minutes. The EBCTs for two full-scale anion exchange plants removing arsenic were 3.7

minutes for Plant A and 5.6 minutes for Plant B (5). 
E-2

 2. The estimated cost of the anion exchange resin was $125/ft3 for all system sizes.

Basis. The November 1999 Technology and Cost Document (2) recommended this cost based

on costs at existing ion exchange facilities and conversations with suppliers. The case study

costs in the AwwaRF Document (4) use a range of $110/ft3 at 1 mgd to $90/ft3 at 30 mgd.

 3. The bed depth assumption ranged from 3 feet to 6 feet depending upon the design flow. The

maximum diameter per column was 12 feet. When the calculated media volume exceeded the

maximum, the treatment media was divided into several beds to maintain a realistic bed

diameter. Thus, larger designs include multiple anion exchange columns, each treating a

portion of the total flow.

Basis. The costs for the activated alumina used this approach (6).

 4. The vessel cost has been sized based on 50% bed expansion during backwash. The vessel

volume was calculated as (1.5)(media volume).

Basis. The Water Model used 50% bed expansion to size the vessel (3). 

 5. The vessel cost is based on the following equation: 

Cost = 63.288 * (Vessel Volume in gallons) ^ 0.679

Basis. Quotes were obtained from two manufacturers for glass-lined carbon steel vessels with

working pressure rating between 75 and 150 psi and meeting American Society of Mechanical

Engineers (ASME) Code Section VIII requirements. Sizes ranging from 140 gallons to 6764

gallons were used to develop the equation. Vessel costs for sixteen different sizes were

regressed against their respective volumes to derive the equation.
E-3

 6. The capital costs include a redundant column to allow the system to operate while the media

is being regenerated in the other column.

Basis. A redundant vessel has been included for each anion exchange column. While the anion

exchange column is being regenerated, the other column can produce drinking water.

Redundant columns were not included at either of the two full-scale arsenic anion exchange

plants (5).

 7. The other equipment cost elements are the brine tank, brine pump and mixer. Costs for the

brine tank were based on 30-day storage. Separate costs were prepared for the two sulfate

ranges based on the number of regenerations per month. 

Basis. Brine tank costs were based on the salt needed to perform regenerations every 1500

bed volumes and 700 bed volumes over a 30 day period. Cost data from RS Means

“Environmental Remediation Cost Data” on plastic storage tanks were used to develop the

brine tank costs (7). Data from this source were also used for the brine pump costs. A

manufacturer cost estimate was used for the brine mixer.

 8. The remaining components of the process equipment include concrete, pipe and valves and

electrical and instrumentation. A cost factor of 25.02% of the manufactured equipment and

activated alumina was used for these other construction cost components.

Basis. The Water Model provided construction cost estimates for a variety of conceptual

designs (3). Construction cost components included process costs such as manufactured

equipment and media, concrete, pipes and valves, electrical and instrumentation as well as

construction costs such as excavation and sitework. The weighted percentages for concrete,

pipe and valves and electrical and instrumentation were expressed as a percentage of the

manufactured equipment costs. Resin costs were included in the manufactured equipment

costs. Concrete was 7.32%, steel was 9.51%, pipe and valves were 3.02%, and electrical

and instrumentation were 5.16% of the equipment costs. The total was 25.02%. 
E-4

 9. The capital costs have been estimated from the total process costs using a factor of 2.5 for

small systems and 3.33 for large systems.

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 (8). The process costs

were assumed to be 40% of the capital costs for small systems. The breakdown of capital

costs for small systems is as follows: 40% process costs, 40% construction costs, and 20%

engineering costs. For large systems, the breakdown of capital cost is: 30% process costs,

40% construction, and 30% engineering.

 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. 

Basis. Data from the 1995 Community Water Systems Survey were used to identify the

percentage of systems without treatment (9). The data from ground water systems is as

follows:

Table E-1.

Percentage of Ground Water Systems with No Treatment: Systems Serving # 3,300 People

Population Category

25 - 100 101 - 500 501 - 1,000 1,001 - 3,300

43% 19% 16% 18%
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Table E-2

Percentage of Ground Water Systems with No Treatment: Systems Serving > 3,300 People

Population Category

3,301 - 10,000 10,001 - 50,000 50,001 - 100,000 100,001 - 1 MIL

13% 1% 11% 0%

Larger ground water systems rely on multiple entry points to supply ground water rather than

one well. Thus, treatment may be present in the system, but only at one of the wells. Since the

analysis is based on no treatment throughout the system, using the percentages for larger

systems could underestimate the need for roads and fences at some entry points. Therefore,

the 18% without treatment from the 1,001 - 3,300 category was applied to the four larger

system categories. The following table summarizes the percentage used for each of the size

categories.

Table E-3

Assumed Percentage of Ground Water Systems with No Treatment

Population Category Percentage of GWS w/o any Treatment

25 - 100 43%

101 - 500 19%

501 - 1,000 16%

1,001 - 3,300 18%

3,301 - 10,000 18%

10,001 - 50,000 18%

50,001 - 100,000 18%

100,001 - 1 MIL 18%
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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 = 118.7 * Design Flow + 118

Basis. The basis for the cost equations for buildings, fence and road are described in Section

4.0 of the Very Small Systems Document (10). The Very Small Systems Document provided

costs as a percentage of the capital costs for design flows from 0.024 to 0.65 mgd. 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:

Process Area (sq ft) = 118.7 * Design Flow + 118

The process area for the conceptual designs in the Water Model (3) were also regressed and

fitted to a linear regression. The conceptual designs covered resin volumes from 4 to 520

cubic feet. The linear regression from this source is:

Process Area (sq ft) = 171.8 * Design Flow + 103

The process area estimates using the Very Small Systems Document equation were selected

to make cost estimates for buildings, road and fence. The process area estimates using this

equation were consistently higher than the Water Model estimates over the entire range of

design flows under consideration. Since there would be a minimum building size,

extrapolating a linear regression below the lowest data point does not create a significant

problem with the intercept.
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 12. The building cost to house the activated alumina system is $45.38/sq ft for design flows at or

below 0.65 mgd. The following equation is used for design flows between 0.65 and 1 mgd:

Building cost = 54.204 * Process Area + 14887

For design flows greater than 1 mgd, the building cost varied based on flow from $100 to $90

per square foot.

Basis. The data from both the Very Small Systems Document and the Water Model were

examined to develop housing costs for the anion exchange system. The housing rate listed in

the Very Small Systems Document was $40/sq ft (10). This rate includes the following: basic

storage building, foundation, electrical wiring, HVAC, and plumbing. The rate is for a prefabricated 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 for activated alumina were used because it was easier to link to design flow

and for consistency purposes with activated alumina costs (3). The costs in the Water Model

were developed in 1983. The ENR Building Cost Index was used to update both sets of costs

to September 1998. The housing rate from the Very Small Systems Document was selected

based on better documentation and the use of more recent data for design flows up to 0.65

mgd. For design flows between 0.65 and 1 mgd, a linear regression of activated alumina

housing cost against process area was used to derive the equation listed above. For design

flows greater than1 mgd, the building cost rate assumptions were the same as those used for

the case studies in the AwwaRF Report for activated alumina (4).

E.2.2 Operation and Maintenance Costs

There are major components to the operation and maintenance (O&M) costs are regeneration

frequency, regeneration dose, and incremental labor.

 13. The critical variable in calculating the regeneration frequency is the run length - the number

of bed volumes until the effluent concentration exceeds a specific concentration. The run
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length when sulfate is at or below 20 mg/L is 1500 bed volumes (BV). The run length when

sulfate is between 20 and 50 mg/L sulfate is 700 BV.

Basis. The November 1999 Technology and Cost Document (2) utilized a breakthrough curve

from a computer program based on equilibrium multicomponent chromatography theory with

constant separation factors (11). The number of bed volumes for 10 percent and 50 percent

breakthrough of influent arsenic were presented as a function of influent sulfate concentration.

The 10 percent breakthrough curve would correspond to 90% removal of arsenic. This

breakthrough curve was presented in Figure 3-37 of the November 1999 Technology and Cost

Document (2). Run length for the entire range was based on the run length at the highest sulfate

concentration. A run length of 1500 BV for sulfate concentrations up to 20 mg/L is based on

the run length at 20 mg/L. A run length of 700 BV for sulfate concentrations in the range of 20

to 50 mg/L is based on the run length at 50 mg/L. This is a very conservative assumption. For

example, a system with 25 mg/L would have a run length of 1300 BV, but is assumed to have

a run length of 700 BV. The actual run length would be nearly double the run length being used

to estimate costs. One other source of conservatism in these estimates. The breakthrough

curve presented in the November 1999 Technology and Cost Document may underestimate

actual run length. The run length observed in pilot testing at Albuquerque was 506 BV when

regeneration was performed with 2.0 equivalent chloride per equivalent resin and downflow

regeneration with 1 molar (M) salt solution (12). The sulfate level was 81.7 mg/L. The

predicted run length would be approximately 420 BV using the graph in the November 1999

Technology and Cost Document (2).

The regeneration frequency can be calculated using the run length (number BV), the size of a

BV (gallons), and the average daily flow. The regeneration frequency is reported in days.

This can also be reported as number of regenerations per week.

 14. The regeneration dose is also a critical element for the O&M costs. The salt dose for

regeneration was 10.2 lb/ft3.

Basis. Regeneration in the Albuquerque study (12) was achieved using a 1 M sodium chloride

solution at a regeneration level of 2 eq chloride/eq resin (10.2 lbs NaCl/ft3 resin).
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 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 comes with a new process.

Adjustment in staffing or shifting of activities are not included in the incremental labor. The

incremental labor for the anion exchange is one hour per week plus three hours per

regeneration.

Basis. The Very Small Systems Document states that fours per day are required for resin

regeneration, routine equipment maintenance and process oversight (10). This section covers

both cation and anion exchange, so this estimate is likely based on daily or more frequent

regeneration for cation exchange softening or nitrate removal. The Water Model (3) lists labor

hours per year based on the frequency of regeneration. The number of hours varies based on

resin volume. For daily regeneration, the range is between 200 and 300 hours per year - less

than one hour per day. One hour per week for routine equipment maintenance and three hours

per regeneration was assumed based on both sources.

 16. The labor rate for small systems was $28/hour. The loaded labor rate for large systems was

$52/hour.

Basis. A rate of $28/hour is recommended as a loaded labor rate for small systems at the EPA

Technology Design Workshop Summary included in Annex A of the Phase I Upgrade

Document (8). A loaded labor rate of $52/hour was selected for large systems based on

licensed operators. For this analysis, a large system was defined as a system with an average

flow greater than 0.36 mgd (corresponds to a design flow of 1 mgd). This corresponds to

populations lower than 3,300 people. The use of a large system rate may overestimate costs

in systems serving between 2,700 and 10,000 people.

The Design Workshop Summary (8) also noted that the Bureau of Labor Statistics (BLS)

publishes monthly reports that include fully burdened labor costs on a regional basis. Other

members of the Panel noted that some States have conducted operator surveys that include

labor rates. The Panel recommended that EPA review and compare both the BLS and State

material to derive representative rates. A preliminary analysis of such material indicates that
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the labor rates for both large and small systems may be very conservative and that actual labor

rates may be much lower.

 17. O&M Costs were calculated using the data on regeneration dose, regeneration frequency,

labor hours per week and adjusted labor rates using the equation in the Very Small Systems

Document.

Basis. The equation allows the critical variables to be used to estimate O&M costs. The

equation produces O&M cost in cents per thousand gallons, so it must be converted to dollars

per year. Run lengths of 700 BV and 1500 BV were included in the November 1999

Technology and Cost Document (2). The O&M costs produced by this costing procedure are

consistently higher than those in the November 1999 Technology and Cost Document even at

the higher flow rates. This is true for both run lengths. Even though the equation in the Very

Small Systems Document listed a domain up to 0.27 mgd, it appears to function over the entire

flow range.

E.3 WASTE DISPOSAL COSTS

 18. The primary waste is the regenerant brine stream. Since arsenic in the brine stream would

likely exceed the toxicity characteristic regulatory level of 5 mg/L, it could be considered

hazardous waste. Any mixture of domestic sewage and other wastes that passes through a

sewer system to a publicly-owned treatment works for treatment is excluded from being solid

waste and thus can’t be hazardous waste. Discharge to a POTW is the only waste disposal

option being considered for anion exchange brine. 

Basis. The proposed rule listed three mechanisms to dispose of the brine stream used for

regeneration. The options were: sanitary sewer, evaporation pond, and chemical

precipitation. Many comments on the proposed rule were concerned about waste streams

being considered hazardous waste. Waste streams containing less than 0.5% solids are

evaluated against the toxicity characteristic directly to determine if the waste is hazardous.
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Arsenic in the regeneration brine will likely exceed 5 mg/L for most systems with arsenic

above 15 Fg/L and sulfate below 50 mg/L. Since the brine stream would likely be considered

hazardous, the evaporation pond and the chemical precipitation options were eliminated from

the decision tree as options for disposal of anion exchange wastes. Discharge to a sanitary

sewer was retained because domestic sewage and any mixture of domestic sewage and other

wastes that pass through a sewer system to a publicly-owned treatment works (POTW) for

treatment is excluded from being considered solid waste (40 CFR 261.4). Domestic sewage

means untreated sanitary wastes that pass through a sewage system. Discharges meeting the

above criteria are excluded from regulation as hazardous waste.

 19. The volume of the waste stream is 7.7 bed volumes (BV).

Basis: The volume of the salt solution used for regeneration is based on an equation in Water

Quality and Treatment (13). The critical variables are the regenerant dose and the molarity

of the salt solution. The brine volume is 2.7 BV. The displacement rinse volume can range

from 2.5 to 5 BV (12). The upper bound of 5 BV was used to calculate waste volume for

O&M costs. The total waste stream is 7.7 BV.

 20. Capital and O&M costs were calculated using the two residuals management document for

discharge to POTW.

Basis: Equations in the two water system byproducts were used for the capital and O&M costs

for discharge to POTW (14, 15). Costs were updated to September 1998 for consistency with

other cost equations.

E.4 REFERENCES

 1. US EPA. Technologies and Costs for Removal of Arsenic from Drinking Water. Prepared

by International Consultants, Inc. and Malcolm Pirnie, Inc. under contract 68-C6-0039 for EPA

OGWDW. April 1999.
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 2. US EPA. Technologies and Costs for Removal of Arsenic from Drinking Water. Prepared

by International Consultants, Inc. and Malcolm Pirnie, Inc. under contract 68-C6-0039 for EPA

OGWDW. November 1999.

 3. US EPA. Estimation of Small System Water Treatment Costs, Prepared by

Culp/Wesner/Culp under contract 68-03-3093 for EPA ORD. 1983.

 

 4. American Water Works Association Research Foundation. Cost Implications of a Lower

Arsenic MCL. Prepared by Frey, M.M. et al for American Water Works Association Research

Foundation. May 5, 2000.

 5. US EPA. Arsenic Removal from Drinking Water by Ion Exchange and Activated Alumina

Plants. Prepared by Battelle under contract 68-C7-0008 for EPA ORD. October 2000.

 6. US EPA. “Basis for Arsenic Activated Alumina Costs for Final Rule.” December 13, 2000.

 7. R.S. Means. Environmental Remediation Cost Data. 1998.

 8. US EPA. Guide for Implementing Phase I Water Treatment Cost Upgrades. Prepared by

Science Applications International Corporation under contract 68-C6-0059 for EPA

OGWDW. September 1998.

 9. US EPA. Geometries and Characteristics of Public Water Systems. Prepared by Science

Applications International Corporation under contract 68-C6-0059 for EPA OGWDW. August

15, 1999.

 10. US EPA. Very Small Systems Best Available Technology Cost Document. Prepared by

Malcolm Pirnie, Inc. for EPA. September 1993.

 11. Clifford, D.A. “Computer Prediction of Arsenic Ion Exchange.” Journal AWWA. April 1995.
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 12. Clifford, D.A. G. Ghurye et al. Final Report: Phases 1 & 2 City of Albuquerque Arsenic

Study Field Studies on Arsenic Removal in Albuquerque, New Mexico using the University

of Houston/EPA Mobile Drinking Water Treatment Research Facility. Houston, TX:

University of Houston, December 1997.

 13. American Water Works Association. Water Quality and Treatment - A Handbook of

Community Water Supplies. Fifth Edition. McGraw-Hill Publishing Company, New York.

1999.

 14. US EPA. Small Water System Byproducts Treatment and Disposal Cost Document.

Prepared by DPRA Incorporated for EPA OGWDW. April 1993.

 15. US EPA. Water System Byproducts Treatment and Disposal Cost Document. Prepared by

DPRA Incorporated for EPA OGWDW. April 1993.
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