EPA 625/1-76-OOla
PROCESS DESIGN MANUAL
FOR
PHOSPHORUS REMOVAL
U.S. ENVIRONMENTAL PROTECTION AGENCY
Technology Transfer
U.S. Environmental Protection Agfc:v-y
Great Lakes National Program Offica
.GLHPO Library
April 1976
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ACKNOWLEDGMENTS
The original edition of this Design Manual (October 1971) was prepared for Technology
Transfer by Black & Veatch, Consulting Engineers. This first revision to the basic text was
prepared for Technology Transfer by the firm of Shimek, Roming, Jacobs and Finklea.
Major EPA contributors were E. F. Earth, J. M. Smith, C. A. Brunner, and J. B. Farrell of
the U.S. EPA Office of Research and Development, Cincinnati, Ohio. Major reviewers were
D. J. Lussier, J. C. Dyer, and R. M. Madancy of Technology Transfer.
Revisions to Chapters 4, 5, 10 and 11 have been made by J. Laughlin of Shimek, Roming,
Jacobs and Finklea in order to update the information in these chapters. Appendix A has
been added to describe an EPA computer model for evaluating phosphorus removal strate-
gies.
NOTICE
The mention of trade names or commercial products in this publication is for illustration
purposes and does not constitute endorsement or recommendation for use by the U.S.
Environmental Protection Agency.
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ABSTRACT
The discharge of phosphorus-containing wastewaters into the surface waters of the
United States has contributed to their over-fertilization and eutrophication. As a result,
efforts are now being made to remove phosphorus from wastewater.
This manual discusses phosphorus removal methods that have been found effective and
practical for use at treatment plants. All the methods included involve chemical
precipitation of the phosphorus and removal of the resultant precipitate. Precipitants
include salts of aluminum and iron, and lime. The practical points of addition are
before the primary settler, in the aerator of an activated sludge plant, before the final
settler, or in a tertiary process.
Included in the discussion of each treatment method is a description of the method,
pilot or full-scale performance data, equipment requirements, design parameters, and
costs. This information should be of value to designers, municipal officials, regulatory
agencies, city planners, and treatment plant operators.
111
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TABLE OF CONTENTS
CHAPTER PAGE
ACKNOWLEDGMENTS ii
ABSTRACT iii
TABLE OF CONTENTS v
LIST OF FIGURES vii
LIST OF TABLES xiii
FOREWORD xv
1 INTRODUCTION
1.1 Purpose 1-1
1.2 Scope 1-1
2 BASIC DESIGN CONSIDERATIONS
2.1 Sources of Phosphorus 2-1
2.2 State Regulatory Agencies 2-2
2.3 State Standards 2-2
2.4 Provision for Future Plant Expansion and More
Stringent Effluent Requirements 2-3
2.5 Flow Equalization 2-3
2.6 References 2-5
3 THEORY OF PHOSPHORUS REMOVAL BY CHEMICAL
PRECIPITATION
3.1 Forms and Measurement of Phosphorus 3-1
3.2 Chemistry of Removal by Precipitation 3-2
3.3 References 3-6
4 PHOSPHORUS REMOVAL BY MINERAL ADDITION
BEFORE THE PRIMARY SETTLER
4.1 General Considerations 4-1
4.2 Phosphorus Removal Data for Alum Addition 4-4
4.3 Phosphorus Removal Data for Iron Addition 4-5
4.4 Overall Performance of Treatment Plants with
Mineral Addition to the Primary Settler 4-8
4.5 Dosage Selection and Correlation 4-10
4.6 Case Histories 4-13
4.7 Costs 4-21
4.8 References 4-22
5 PHOSPHORUS REMOVAL BY LIME ADDITION BEFORE
THE PRIMARY SETTLER
5.1 Introduction 5-1
5.2 System Layout and Operation 5-2
5.3 Case Histories 5-5
5.4 Other Experience 5-15
5.5 Sludge 5-23
5.6 Costs 5-30
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TABLE OF CONTENTS - Continued
CHAPTER PAGE
5 PHOSPHORUS REMOVAL BY LIME ADDITION BEFORE
THE PRIMARY SETTLER - Cont'd.
5.7 References 5-31
6 PHOSPHORUS REMOVAL IN TRICKLING FILTERS BY
MINERAL ADDITION
6.1 Pre-Design Decisions 6-1
6.2 Process Options 6-1
6.3 Performance Data Using Aluminum Salts 6-2
6.4 Performance Data Using Iron Salts 6-8
6.5 Choice of Chemical Addition 6-8
6.6 Nature and Role of Chemicals Involved 6-9
6.7 Dosage Selection and Control 6-10
6.8 Iron Leakage 6-13
6.9 Sampling and Analysis 6-13
6.10 References 6-16
7 PHOSPHORUS REMOVAL IN ACTIVATED SLUDGE
PLANTS BY MINERAL ADDITION
7.1 Description of Process 7-1
7.2 Mineral Selection and Addition 7-3
7.3 Performance and Optimization 7-9
7.4 Process Design Examples 7-18
7.5 References 7-26
8 PHOSPHORUS REMOVAL BY LIME TREATMENT OF
SECONDARY EFFLUENT
8.1 Description of Process 8-1
8.2 Typical Performance Data 8-2
8.3 Criteria for Selection of Process 8-9
8.4 Description of and Criteria for Choice of Equipment 8-10
8.5 Capital and Operating Costs 8-19
8.6 References 8-21
9 PHOSPHORUS REMOVAL BY MINERAL ADDITION TO
SECONDARY EFFLUENT
9.1 Description of Process 9-1
9.2 Summary of Design Information 9-1
9.3 Laboratory and Pilot Studies 9-2
9.4 References 9-7
10 STORAGE AND FEEDING OF CHEMICALS
10.1 Aluminum Compounds 10-1
10.2 Iron Compounds 10-18
10.3 Lime 10-27
10.4 Other Compounds for pH Adjustment 10-33
VI
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CHAPTER
10
11
TABLE OF CONTENTS - Continued
STORAGE AND FEEDING OF CHEMICALS - Cont'd.
10.5 Carbon Dioxide
10.6 Polymers
10.7 Chemical Suppliers
10.8 Chemical Feeders
10.9 Additional Reading
10.JO References
SLUDGE HANDLING AND DISPOSAL
11.1 Introduction
11.2 Generalized Assessment of Filtration Rates
11.3 Quantity of Phosphorus Removal Sludges
11.4 Dewatering at Specific Locations
11.5 References
PAGE
1040
1043
1047
10-66
10-75
10-77
11-1
11-2
11-6
11-29
11-38
Appendix
A
B
LIST OF APPENDIXES
EVALUATING PHOSPHORUS REMOVAL STRATEGIES
METRIC CONVERSION CHART
A-l
B-l
Vll
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LIST OF FIGURES
Figure No. Page
4-1 Lebanon, Ohio Primary Treatment and Chemical Feeding
Schematic 4-2
4-2 Schematic of Pre-Treatment of Plant Flow at Benton Harbor-
St. Joseph Wastewater Disposal Plant 4-3
4-3 Soluble Phosphorus Removal by Ferric Chloride Addition 4-11
4-4 Diurnal Variation of Ortho Phosphorus 4-12
4-5 Treatment Facilities at Pontiac, Michigan 4-15
4-6 Treatment Facilities at Trenton, Michigan 4-19
5-1 Typical Facilities for Low-Lime Addition in Primary Treatment 5-4
5-2 Treatment Plant at Newmarket, Ontario 5-6
5-3 System Layout at Contra Costa County 5-10
5-4 Treatment System at Rochester, N.Y. 5-17
5-5 Typical Effect of pH on Phosphorus Removal 5-20
5-6 Phosphorus Removal Process at Seneca Falls, N.Y. 5-24
5-7 Proposed Sludge Facility at Contra Costa County 5-27
5-8 Plant Scale Solids Testing System at Contra Costa 5-28
5-9 Sludge Classification Curves at Contra Costa 5-29
6-1 Richardson, Texas Treatment Plant Facilities 6-3
6-2 Chapel Hill, N.C. Treatment Plant Facilities 6-6
6-3 Phosphorus Removal by Iron Addition (Richardson, Texas) 6-12
6-4 Iron Leakage During Phosphorus Removal 6-14
7-1 Phosphorus Removal Capabilities of Activated Sludge Systems
Receiving Settled Domestic Wastewater 7-2
7-2 Total Soluble Phosphorus Residual versus Metal Ion Dose 7-4
7-3 Influence of Point of Addition on Phosphorus Removal with
Aluminum 7-5
7-4 Effect of pH on Phosphorus Removal from a Final Effluent 7-7
IX
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LIST OF FIGURES - Continued
Figure No. Page
7-5 Comparison of the Effectiveness of Alum and Sodium
Aluminate on Phosphorus Removal 7-11
7-6 Chemical Phosphorus Removal Envelope: High Rate
Activated Sludge System 7-15
7-7 0.2 mgd Biological Nitrification-Denitrification of Pilot Plant 7-16
7-8 Effect of Aluminum Addition with pH Adjustment on Total
Soluble Phosphorus Residual 7-18
7-9 Physical Characteristics of High Rate Activated Sludge with
Mineral Addition (8, 16) 7-22
8-1 Single Stage Lime Treatment System 8-3
8-2 Two Stage Lime Treatment System 8-4
8-3 Effect of pH on Phosphorus Concentration of Effluent from
Filters Following Lime Clarifier 8-7
8-4 Typical Upflow Clarifier 8-11
10-1 Typical Dry Feed System 10-4
10-2 Viscosity of Alum Solutions 10-8
10-3 Alternative Liquid Feed Systems for Overhead Storage 10-10
10-4 Alternative Liquid Feed Systems for Ground Storage 10-11
10-5 Freezing Point Curve for Commercial Ferric Chloride Solutions 10-19
10-6 Viscosity vs Composition of Ferric Chloride Solutions at
Various Temperatures 10-20
10-7 Typical Lime Feed System 10-29
10-8 Viscosity of Soda Ash Solution 10-34
10-9 Viscosity of Caustic Soda Solution 10-36
10-10 Typical Caustic Soda Feed System 10-38
10-11 Manual Dry Polymer Feed System 10-44
10-12 Automatic Dry Polymer Feed System 10-46
10-13 Plunger Type Metering Pump 10-68
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LIST OF FIGURES - Continued
Figure No. Page
10-14 Diaphragm Type Metering Pump 10-69
10-15 Screw Feeder 10-71
10-16 Positive Displacement Solid Feeder-Rotary 10-72
11-1 Effect of pH on Total Residual Phosphorus 11-21
11-2 Ratio of Lime Dose to Alkalinity (Both MG/L) for a
Prescribed Final pH 11-22
11-3 Precipitation of Calcium by Lime Addition 11-25
11-4 Precipitation of Magnesium by Lime Addition 11-26
11-5 Effect of Ferric Chloride on Sludge Dewatering 11-36
XI
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LIST OF TABLES
Table No. Page
4-1 Phosphorus Removal Operating Data for Mineral Addition Before
Primary Clarifier 4-6
4-2 Effectiveness of Primary and Secondary Treatment with and
without Mineral Addition for Phosphorus Removal 4-9
4-3 Design Data for Pontiac, Michigan 4-14
4-4 Operating Data for Pontiac, Michigan 4-16
4-5 Design Data for Trenton, Michigan 4-18
4-6 Operating Data for Trenton, Michigan 4-20
5-1 Design and Operating Data for Newmarket, Ontario 5-5
5-2 Operating Results from Newmarket, Ontario 5-8
5-3 Construction Costs at Newmarket, Ontario 5-9
54 Effluent Requirements for CCCSD Treatment Facility 5-9
5-5 Basic Data for CCCSD 5-11
5-6 Lime Treatment Results at CCCSD (12) 5-13
5-7 Lime and Metal Treatment Results at CCCSD (12) 5-13
5-8 Design Data for Rochester, N.Y. 5-16
5-9 Results of Phosphorus Trial at Seneca Fails, N.Y. 5-25
5-10 Typical Capital Costs for Lime Treatment in Primary Facilities 5-30
5-11 Typical Total Costs for Lime Treatment in Primary Facilities 5-31
6-1 Operating Data for Richardson, Texas 6-4
6-2 Operating Data for Chapel Hill, North Carolina 6-7
6-3 Iron Addition for Phosphorus Removal at Wyoming, Michigan 6-9
6-4 Desirable Laboratory Analyses 6-15
7-1 Metal Salt Comparison 7-8
7-2 Comparison of Use of Alum and Sodium Aluminate at
Pennsylvania State University 7-12
xin
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LIST OF TABLES - Continued
Table No. Page
7-3 Comparison of Alum Addition to the Aerator and Conven-
tional Activated Sludge at Pennsylvania State University
for Flows Not Exceeding the Design Capacity of the Plant 7-13
7-4 Combined Chemical-High Rate Activated Sludge System
Performance 7-14
7-5 Summary of Design Calculations 7-24
8-1 Lime Requirements 8-8
8-2 Capital Cost of Lime Treating Facilities 8-19
8-3 Total Cost for Lime Treatment of Wastewater 8-19
9-1 1966 Estimated Costs for Tertiary Phosphate Removal with
Alum 9-2
9-2 Nassau County Performance Data 9-3
9-3 Nassau County Cost Estimates 9-3
9-4 Removal of Total Orthophosphate by Coagulation-Clarification
in the Dallas Demonstration Plant 9-4
10-1 Available Grades of Dry Alum 10-1
10-2 Solubility of Alum at Various Temperatures 10-2
10-3 Crystallization Temperatures of Liquid Alum 10-7
10-4 Suppliers of Aluminum Chloride 10-16
10-5 Cost of Liquid Alum Feed Facilities 10-17
10-6 Properties of Ferric Sulfate 10-23
10-7 Cost of Liquid Ferric Chloride Feed Facilities 10-26
10-8 Cost of Lime Feed Facilities 10-32
10-9 Materials Suitable for Caustic Soda Service 10-39
10-10 Carbon Dioxide Yields of Common Fuels 10-40
10-11 Cost of Polymer Feed Facilities 10-47
xiv
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LIST OF TABLES - Continued
Table No. Page
10-12 Geographical Summary of Suppliers of Metal Salts and Lime 10-4-8
10-13 Suppliers of Metal Salts and Lime 10-49
10-14 Polymer Sources and Trade Names 10-59
10-15 Soda Ash Manufacturers 10-65
10-16 Types of Chemical Feeders 10-74
11-1 Hypothetical Compilation of Dewatering Rates 11-2
11-2 Design Factors for Filtration of Conventional Sludges 11-4
11-3 Design Factors for Filtration of Conventional Plus Metal Sludges 11-5
11-4 Calculation of Chemical Sludge Production 11-7
11-5 Removal of Dissolved Materials During Chemical Addition 11-11
11-6 Calculation of Total Sludge Mass with Metal Addition 11-14
11-7 Information Needed to Calculate Lime Sludge Mass 11-23
11-8 Calculation of Lime Sludge Mass 11-27
xv
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FOREWORD
The formation of the United States Environmental Protection Agency marked a new era
of environmental awareness in America. This Agency's goals are national in scope and
encompass broad responsibility in the area of air and water pollution, solid wastes,
pesticides, and radiation. A vital part of EPA's national water pollution control effort is
the constant development and dissemination of new technology for wastewater treatment.
It is now clear that only the most effective design and operation of wastewater treatment
facilities, using the latest available techniques, will be adequate to meet the future water
quality objectives and to ensure continued protection of the nation's waters. It is
essential that this new technology be incorporated into the contemporary design of waste
treatment facilities to achieve maximum benefit of our pollution control expenditures.
The purpose of this manual is to provide the engineering community and related industry
a new source of information to be used in the planning, design and operation of present
and future wastewater treatment facilities. It is recognized that there are a number of
design manuals, manuals of standard practice, and design guidelines currently available in
the field that adequately describe and interpret current engineering practices as related to
traditional plant design. It is the intent of this manual to supplement this existing body
of knowledge by describing new treatment methods, and by discussing the application of
new techniques for more effectively removing a broad spectrum of contaminants from
wastewater.
Much of the information presented is based on the evaluation and operation of pilot,
demonstration and full-scale plants. The design criteria thus generated represent typical
values. These values should be used as a guide and should be tempered with sound
engineering judgment based on a complete analysis of the specific application.
This manual is one of several available through the Technology Transfer Office of EPA to
describe recent technological advances and new information. This particular manual was
initially issued in October of 1971 and this edition represents the first revision to the
basic text. Future editions will be issued as warranted by advancing state-of-the-art to
include new data as it becomes available, and to revise design criteria as additional
full-scale operational information is generated.
xvn
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Chapter 1
INTRODUCTION
1.1 Purpose
The technology for removal of phosphorus from wastewater has developed rapidly in
the last few years. The need for practical phosphorus removal procedures is a result of
the over-fertilization and eutrophication of the Country's surface waters. This further
resulted in the establishment of state water quality standards that limit the
concentration of phosphorus in receiving waters. As an example, states that have
streams tributary to the Great Lakes have established wastewater effluent phosphorus
limits or require a fixed percentage reduction of phosphorus. Several other states have
also established phosphorus standards.
The Environmental Protection Agency has sponsored many research and demonstration
studies at several cities in the past few years to advance the knowledge of phosphorus
removal. Local and state governments and private industries have also contributed to
this work. This manual is intended to summarize process design information for the
best developed removal methods that have resulted from this governmental and private
effort.
1.2 Scope
This manual discusses a number of phosphorus removal methods that involve chemical
precipitation. Phosphorus removal obtainable by biological activity alone is not in-
cluded. Treatment methods in which phosphorus removal occurs, but is not a principal
objective, are also omitted. The latter group of processes includes ion exchange, reverse
osmosis and other demineralization treatments which at present are more closely
associated with wastewater renovation and reuse than with pollution control. These will
be included in updated versions of the manual when appropriate. One application of a
precipitation method that has been omitted and which is of growing interest is
complete lime treatment of raw sewage. This application will be included when the
manual is updated.
The information included on chemical methods of phosphorus removal was obtained
from the available literature, progress reports of demonstration studies, and private
communications with investigators actively working in the field. Design guidelines have
been developed from these sources.
The information contained in this manual is oriented toward design methods and
operating procedures for removal of phosphorus from wastewater to comply with state
water quality standards. It is recognized that present state standards may eventually be
superseded with more stringent requirements for phosphorus removal. Such revisions in
state standards must be considered in design of treatment facilities to minimize future
modifications to treatment plant structures.
1 - 1
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Engineers are often faced with circumstances that require the design of treatment
facilities with a limited amount of data, testing, and time. Such circumstances require
a more conservative design to allow for contingencies. Since most phosphorus removal
methods are of recent development, they fall into this category. Conservative values are
given, therefore, when stating design parameters.
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Chapter 2
BASIC DESIGN CONSIDERATIONS
2.1 Sources of Phosphorus
Domestic wastewater normally has a substantial concentration of phosphorus with the
primary sources being a result of man's activities in the home. Human wastes such as
feces, urine, and waste food disposal account for about 30 to 50% of the phosphorus
in domestic wastewater (1). Detergents containing phosphate builders and used princi-
pally for laundering of clothes account for the remainder of phosphorus, or about 50
to 70%. Other sources of phosphorus may cause deviation from these percentages. For
example, where sodium hexametaphosphate or other phosphorus compounds are used
as corrosion and scale control chemicals in water supplies, the phosphorus added will
be present in the same concentration, although not necessarily in the same form, as in
the water supply. This source can account for 2 to 20% of the total phosphorus
present in the wastewater.
Industrial wastewater discharges to the sanitary sewer system may either dilute or
increase the concentration of phosphorus present in the combined wastewater. For
example, wastewaters from pulp and paper mills may be phosphorus deficient and
therefore dilution will occur. On the other hand, wastewaters discharged from some
industries, such as potato processing plants, may contain high concentrations of phos-
phorus and consequently may increase the phosphorus concentration of the combined
wastewater.
The quantity of phosphorus resulting from human excretions reportedly ranges from
0.5 to 2.3 Ib per capita per year (2). The mean annual excretion is estimated to be
about 1.2 Ib per capita. The mean annual contribution of phosphorus from synthetic
detergents with phosphate builders is estimated to be about 2.3 Ib per capita, at
present. Thus, exclusive of industrial wastes and other phosphorus sources, such as
water softening or sequestering agents, the domestic phosphorus contribution to waste-
water is about 3.5 Ib per capita per year.
On this basis, the wastewater phosphorus concentration may be estimated for a
community where there are no phosphorus data available. Assuming a population of
5,000, the total phosphorus discharged equals 5,000x3.5= 17,500 Ib/year, or
48 Ib/day. If a wastewater flow of 100 gal. per capita per day is assumed, the total
flow is 0.5 million gal. per day (mgd) and the phosphorus concentration in the
wastewater is as follows:
48
0.5 x 8.34
= 11.5 mg/1 (asP)
The average total phosphorus concentration in domestic raw wastewater is found to be
about 10 mg/1 expressed as elemental phosphorus (P).
The above information is provided to serve as a rough guide to engineers and not as a
basis of design. Sampling of the wastewater and analyses for phosphorus are
2 - 1
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recommended in all cases. It may also be necessary to survey the industries that will
be discharging wastewater to the sanitary sewer system, to determine their influence on
the concentration of phosphorus in the combined wastewater.
Among industrial sources of phosphorus may be potato processing wastes (3), fertilizer
manufacturing wastes (4), animal feedlot wastes (5), certain metal finishing wastes (1,6),
flour processing wastes (7), dairy wastes (8, 9), commerical laundry wastes (10), and
slaughterhouse wastes (11). The amount of phosphorus may be quite variable depending
on the specific industrial plant.
2.2 State Regulatory Agencies
Each state has at least one organization responsible for water pollution prevention,
control, and abatement activities. In about half of the 50 states this responsibility is
divided between two state agencies. Usually one of these is the state health department
and the other an organization charged specifically by statute with the conduct of the
state's water pollution control program. Although there has been an intensification of
Federal activities in water pollution control over the past twenty years, much of the
responsibility for program operations rests with the states. Invariably the appropriate
state agencies should be consulted in connection with the planning and conduct of
water pollution control programs at cities and other communities. The Environmental
Protection Agency regional office personnel work closely with state agencies and can
give advice regarding the proper state agency contacts.
2.3 State Standards
As of June 1971, sixteen states had adopted wastewater effluent phosphorus standards.
In general, these standards have taken the form of an effluent concentration limit or a
requirement for a specified percentage reduction in the phosphorus concentration in
the raw wastewater. In most cases, effluent concentration limits range from 0.1 to
2.0 mg/1 as P, with many established at 1.0 mg/1. Percentage reduction requirements
range from 80 to 95%.
It should be noted that neither an effluent nor a percentage-reduction standard
actually limits the phosphorus load in terms of pounds of phosphorus discharged per
day. Load (Ib/day) is a direct function of both effluent P concentration and daily
effluent volume. Assuming the effluent concentration standard is being met, an increase
in the daily flow of effluent will produce a proportionate increase in the mass of
phosphorus discharged daily to the receiving waters. If the phosphorus load then
becomes intolerable, adjustment downward of the effluent concentration standard will
be necessary. Similarly, adjustment upward of a percentage-reduction standard will be
required if the phosphorus load becomes objectionable as a result of increased waste-
water flow or increased P concentration in the raw wastewater or both.
Effluent concentration standards can be derived from receiving water standards as
indicated in the following example.
Receiving Water
Phosphorus concentration upstream 0.01 mg/1
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Receiving Water - (continued)
Maximum allowable phosphorus concentration 0.05 mg/1
Low flow 140 ft3/sec (90 mgd)
Wastewater
Average daily flow 4.8 mgd
Effluent concentration standard X
Mass Balance
(0.01)(8.34)(90) + (X)(8.34)(4.8) = (0.05)(8.34)(90+4.8)
X = 0.8 mg/1 P
Implicit in the above calculation is the assumption that the effluent is mixed thorough-
ly with the waters of the receiving stream. The low flow selected is not necessarily the
low flow of record. It is usually the low flow averaged over a period of 10 days or
less and having a return period of once in 10 years.
2.4 Provision for Future Plant Expansion
and More Stringent Effluent Requirements
The plant site is an important consideration in the design of treatment works. It is
highly desirable that sufficient site area be obtained to permit initial construction of
the required system without crowding and to provide space for future expansion.
Treatment systems are likely to become larger and more complex in the future rather
than less so, and failure to obtain ample site area initially is almost certain to create
costly problems later.
Reduction of the allowable effluent phosphorus concentration may be necessary in
some states as increases in wastewater flow or phosphorus concentration of the raw
wastewater occur. In some instances, merely increasing the chemical dosage may be the
only necessary change. On the other hand, it may be necessary to add an additional
treatment sequence such as filtration to the existing facilities. In nearly every case,
planning for plant expansion or more stringent effluent requirements will result in
economy. Such planning should include selection of a plant layout that allows space
for addition of facilities without removing existing processes from service.
Other provisions that should be considered are space within chemical buildings for
additional coagulant storage and feeding equipment, space for polymer storage and feed
equipment, and adequate space for handling the additional solids generated by phos-
phorus removal processes.
2.5 Flow Equalization
The cyclic nature of wastewater flows, in terms of volume and strength, is well
established. While the concept of flow equalization has been employed in the field of
water supply, and in the treatment of some industrial wastes, it has not been widely
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accepted in the municipal pollution control field. Anticipated problems with solids
settling, odor, and septicity can be cited as the major factors limiting its use.
The advent of stricter stream standards, which sometimes require removal of contami-
nants during peak flows, the elimination of plant bypassing, and the increased removal
efficiencies that are possible when biological or chemical treatment processes are
operated at near steady-state conditions, have all favored the increased use of flow-
smoothing, and flow equalization devices ahead of or integral with the design of major
treatment works.
There are two major objectives in the design of flow equalization basins. The first of
these is simply to dampen the diurnal flow variation that normally exists in typical
municipal wastewater collection systems to achieve a constant or nearly constant flow
rate through the downstream treatment processes. In this type of system, little
consideration is given to controlling the concentration changes that take place during
storage. The major design factors are supplying enough air to keep the basin aerobic
and to provide adequate mixing to prevent solids deposition.
The second objective of flow equalization is to provide the capacity to distribute
shock loads of toxic or process inhibiting substances over a reasonable period of time
so as to prevent system failure and to minimize the periodic discharge of harmful
contaminants to the receiving stream or surface impoundment. Time-dependent concen-
tration profiles and flow-through curves are normally used to analyze the flow charac-
teristics of these systems and to determine the effect of tank geometry, placement of
effluent weirs and mixing regime on changes in contaminant concentrations through
the basin.
In all cases the added costs of flow equalization must be measured against the possible
reduction in downstream process costs and the increased efficiencies that can be
achieved by operating these processes under constant loading conditions.
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2.6 References
1. "Phosphates in Detergents and the Eutrophication of America's Waters", Hearings
before a Subcommittee of the Committee on Government Operations, House of
Representatives, 91st Congress, December 15-16, 1969, U.S. Government Printing
Office, Washington (1970).
2. Task Group Report, "Sources of Nitrogen and Phosphorus in Water Supplies",
JAWWA, 59:3, p 344 (1967).
3. Mackenthun, K. M., "The Phosphorus Problem", JAWWA, 60:9, p 1047 (1968).
4. Nemerow, N. L., Theories and Practices of Industrial Waste Treatment, Addison -
Wesley, Reading, Mass. (1963).
5. Relationship of Agriculture to Soil and Water Pollution, Cornell University Confer-
ence on Agricultural Waste Management, Rochester, N.Y., Jan. 19-21, 1970.
6. Anderson, J. S., and lobst, E. H., Jr., "Case History of Wastewater Treatment in a
General Electric Appliance Plant", JWPCF, 40:10, p 1786 (1968).
7. Dickerson, B. W., and Farrell, P. J., "Laboratory and Pilot Plant Studies on
Phosphate Removal from Industrial Wastewater", JWPCF, 41:1, p 56 (1969).
8. McKee, F. J., "Dairy Waste Disposal by Spray Irrigation", Sew. & Ind. Wastes,
29:2, p 157 (1957).
9. Lawton, G. W., et al., "Spray Irrigation of Dairy Wastes", Sew. & Ind. Wastes,
31:8, p 923 (1959).
10. Flynn, J. M., and Andres, B., "Launderette Waste Treatment Processes", JWPCF,
35:6, p783 (1963).
11. Wymore, A. H., and White, J. E., "Treatment of Slaughterhouse Waste Using Anaer-
obic and Aerated Lagoons", Wat. & Sew. Works, 115:10, p 492 (1968).
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Chapter 3
THEORY OF PHOSPHORUS REMOVAL
BY CHEMICAL PRECIPITATION
3.1 Forms and Measurement of Phosphorus
Phosphorus is found in wastewater in three principal forms: orthophosphate ion,
polyphosphates or condensed phosphates, and organic phosphorus compounds. Ortho-
phosphate was in the past often considered to be PO,^ and was reported in this way.
Actually, there are a number of forms of orthophosphate in equilibrium, with the
predominant form changing as pH changes. At the usual pH of municipal wastewater,
the predominant form is HPO^". With the present practice of reporting phosphorus as
P, one does not usually have to be concerned with the actual form. It is well to
remember, however, that the predominant form does change with pH. The removal of
phosphorus with lime results primarily from pH increase. Polyphosphates can be looked
upon as polymers of phosphoric acid from which water has been removed. Materials of
various molecular weight are in widespread use. Complete hydrolysis results in forma-
tion of orthophosphate. The chemistry of organic phosphorus compounds is complica-
ted. Their decomposition also leads to orthophosphate.
In raw sewage there are substantial amounts of all three principal phosphorus forms.
During biological treatment, significant changes take place. As organic materials are
decomposed, their phosphorus content is converted to orthophosphate. On the other
hand, inorganic phosphates are utilized in forming biological floe. The polyphosphates
are for the most part converted to orthophosphate. The result is that in a well treated
secondary effluent, a large fraction of the phosphorus is present as orthophosphate.
From the standpoint of removal by precipitation, this is fortunate since the orthophos-
phate is the easiest form to precipitate.
Because phosphorus will be present in a variety of forms, both organic and inorganic,
the only satisfactory measure of treatment plant removal efficiency must be based on
total phosphorus entering the plant in the raw wastewater and total phosphorus
discharged in the plant effluent. Inasmuch as all colorimetric tests for phosphorus
depend upon the formation of an orthophosphate color complex, it is essential in the
determination of total phosphorus, to convert any polyphosphates (condensed phos-
phates) and organic phosphorus compounds present in the sample to the orthophos-
phate form. This conversion is accomplished by acid digestion of the sample in the
presence of a strong oxidizing agent. Standard Methods, 13th edition (1), describes
several procedures for determining total phosphorus as well as the various forms of
phosphorus in wastewater. A suitable procedure for measuring total phosphorus, in-
cluding the preliminary digestion stage, may be selected from this reference. An
automated method is described in FWPCA Methods for Chemical Analysis of Water
and Wastes (2).
3 - 1
-------
3.2 Chemistry of Removal by Precipitation
The materials found practical for phosphorus precipitation include the ionic forms of
aluminum, iron, and calcium. The calcium is added as lime and the hydroxyl ions also
have a role in phosphorus removal. The chemistry of phosphorus removal by inter-
action with metallic ions is complex. In the interests of brevity and simplicity, it is
assumed in the following discussion that phosphorus reacts as the orthophosphate
1}
form, PO4 , and that the reaction products are as indicated. The polyphosphates and
organic phosphorus are also removed, probably by a combination of more complex
reactions and sorption on floe particles. The reactions presented and the associated
computations are principally for illustrative purposes and do not purport to represent
the true reactants, products and reaction mechanisms. Calculated requirements for
phosphorus precipitants must be viewed as no more than crude estimates subject to
large variations in actual practice because the forms of phosphorus and the precipitates
may be substantially different from those assumed.
3.2.1 Aluminum Compounds as Phosphorus Precipitants
Aluminum ions can combine with phosphate ions to form aluminum phosphate as
follows:
A13+ + PO43' » A1PO4
The above equation indicates that the mole ratio for A1:PO4 is 1:1. Inasmuch as the
mole ratio of P:PO4 is also 1:1, the mole ratio for A1:P is 1:1 or 4f= 1 when both
aluminum and phosphorus are expressed in terms of gram-moles or Ib-moles. On a
weight, rather than a mole basis, this means that 27 Ib of Al will react with 95 Ib of
PO4 to form 122 Ib of A1PO4. Since each 95 Ib (1 Ib-mole) of PO4 contains 31 Ib
(1 Ib-mole) of P, the weight relationship between Al and P is 27 Ib of Al to 31 Ib of P
or 0.87 for this reaction.
The principal source of aluminum for use in phosphorus precipitation is "alum", a
hydrated aluminum sulfate, having the approximate formula A^CSC^)^ • HI^O (mol-
ecular weight of 594). The chemical, which is also known as "filter alum" or
"papermaker's alum", averages about 17% soluble aluminum expressed as A^O-^ or
about 9.1% expressed as AL Its reaction with PO4 may be written as follows:
A12(S04)3 • 14H20 + 2P043-—*2A1P04 + 3SO42' + 14H2O
The sulfate remains in solution as SO4 . The above reaction indicates that 1 Ib-mole
of alum (594 Ib) will react with 2 Ib-moles (1901b) of PO43" containing 62 Ib
phosphorus to form 2 Ib-moles (244 Ib) of A1PO4- The weight ratio of alum to
phosphorus is, therefore, 9.6:1. The alum requirement per pound of phosphorus may
also be derived from the A1:P mole ratio as follows:
Mole ratio A1:P = 1:1
Therefore weight ratio A1:P = 27:31 = 0.87:1
Alum contains 9.1% Al
0 87
Therefore, alum required per Ib of P = ' = 9.6 Ib
3 - 2
-------
Stumm and Morgan (3), state that the solubility of A1PO4 is pH dependent and varies
as follows:
Approximate
pH Solubility (mg/1)
5 0.03
6 0.01 (minimum solubility)
7 0.3
The optimum pH for removal of phosphorus probably lies in the range of 5.5 to 6.5,
although some removal occurs above pH 6.5.
Addition of alum will lower the pH of wastewater because of neutralization of
alkalinity and release of carbon dioxide. The extent of pH reduction will depend
principally on the alkalinity of the wastewater. The higher the alkalinity, the less is
the reduction in pH for a given alum dosage. Most wastewaters contain sufficient
alkalinity so that even large alum dosages will not lower the pH below about 6.0 to
6.5. In exceptional cases of low wastewater alkalinity, pH reduction may be so great
that addition of an alkaline substance, such as sodium hydroxide, soda ash, or lime will
be required.
Adjustment of the pH downward can be accomplished by the addition of sulfuric acid
but this complicates the treatment process and it may be preferable simply to use a
higher alum dosage.
Bench, pilot, and full-scale studies have shown that considerably higher than stoi-
chiometric quantities of alum usually are necessary to meet phosphorus removal
objectives. A competing reaction, responsible for the pH reduction mentioned above, at
least partially accounts for the excess alum requirement. It occurs as follows:
A12(SO4)3 • 14H2O + 6HCO3'—»2A1(OH)3 + 6CG>2 + 14H2O + 3SO42'
The following ratios of alum (9.1%A1) to phosphorus are believed reasonably repre-
sentative for alum treatment of municipal wastewater (4).
D D . ,. A1:P Alum:P
P Reduction —
Required Mole Ratio Weight Ratio Weight Ratio
75% 1.38:1 1.2:1 13:1
85% 1.72:1 1.5:1 16:1
95% 2.3:1 2.0:1 22:1
To achieve 85% P removal from a wastewater containing 11 mg/1 of P the alum dosage
needed would be (16)(11)= 176 mg/1 or 1470 lb/106 gal.
Sodium aluminate will also serve as a source of aluminum for the precipitation of
phosphorus. The chemical formula for sodium aluminate is Na2Al2O4 or NaAlO2.
3 - 3
-------
One commercial form is the granular trihydrate, which may be written Na2O • AUOo • 3H2O
and which contains about 46% A12O3 or 24% Al.
The sodium aluminate - phosphate reaction may be represented as follows:
Na2O.Al2O3 + 2P043' + 4H2O > 2A1PO4 + 2NaOH + 6OH~
In contrast to alum, which reduces pH, a rise in pH may be expected upon addition
of sodium aluminate to wastewater.
The above reaction indicates an A1:P mole ratio of 1:1, an A1:P weight ratio of
0.87:1, and a sodium aluminate to P weight ratio of about 3.6:1.
As with the use of alum, mole and weight ratios somewhat in excess of the theoretical
must be anticipated in practice.
3.2.2 Iron Compounds as Phosphorus Precipitants
Both ferrous (Fe2+) and ferric (Fe3+) ions can be used in the precipitation of
Q I
phosphorus. With Fe° a reaction can be written similar to that shown earlier for
precipitation of aluminum phosphate. A 1:1 mole ratio of Fe:PO4 results. Since the
molecular weight of iron is 55.85, the weight ratio of Fe:P is 1.8:1. Just as in the case
of aluminum, a larger amount of iron is required in actual situations than the
chemistry of the reaction predicts. With Fe2+ the situation is more complicated and
not fully understood. Ferrous phosphate can be formed. The mole ratio of Fe:PC>4
would be 3:2. Experimental results indicate, however, that when Fe2+ is used, the mole
ratio of Fe:P will be essentially the same as when Fe3+ is used.
A number of iron salts are available for use in phosphorus precipitation. These include
ferrous sulfate, ferric sulfate, ferric chloride, and pickle liquor. Additional discussion of
these is given later. All will lower the pH of wastewater because of neutralization of
alkalinity. Pickle liquor contains substantial amounts of free sulfuric acid or hydro-
chloric acid.
Iron salts are most effective for phosphorus removal at certain pH values. For Fe3+
the optimum pH range is 4.5 to 5.0. This is an unrealistically low pH, not attained in
most wastewaters. Significant removal of phosphorus can be attained at higher pH. For
0-4-
Fe^ the optimum pH is about 8 and good phosphorus removal can be obtained
between 7 and 8. The acidity of ferrous salts, and especially pickle liquor, necessitates
addition of lime or sodium hydroxide for good results. Where the water is aerated
following Fe^+ addition, the use of a base may not be necessary.
3.2.3 Lime as a Phosphorus Precipitant
Calcium ion reacts with phosphate ion in the presence of hydroxyl ion to form
hydroxyapatite. This material has a variable composition, but an approximate equation
for its formation can be written as follows, assuming in this case that the phosphate
present is HPO4 .
3 HPO42' + 5Ca2+ + 4OH~ > Ca5(OH)(PO4)3 + 3 H2O
3 - 4
-------
The reaction is pH dependent. The solubility of hydroxyapatite is so low, however,
that even at a pH as low as 9.0, a large fraction of the phosphorus can be removed.
In lime treatment of wastewater, the operating pH may be predicated on the ability to
obtain good suspended solids removal rather than on phosphorus removal.
Although it is possible to calculate an approximate lime dose for phosphorus removal,
this is generally not necessary. In contrast to iron and aluminum salts, the lime dose is
largely determined by other reactions that take place when the pH of wastewater is
raised. Some of these reactions are discussed later. Only in waters of very low
bicarbonate alkalinity would the phosphate precipitation reaction consume a large
fraction of the lime added.
3 - 5
-------
3.3 References
1. Standard Methods for the Examination of Water and Waste-water, 13th Edition,
American Public Health Assoc., New York (1971).
2. FWPCA Methods for Chemical Analysis of Water and Wastes, Federal Water
Pollution Control Administration, U.S. Department of the Interior (Nov, 1969).
3. Stumm, W. and Morgan, J. J., Aquatic Chemistry, p 521 Wiley-Interscience,
John Wiley & Sons, Inc., New York (1970).
4. Allied Chemical Corporation, Industrial Chemicals Division, "Use of Aluminum
Sulfate for Phosphorus Reduction in Wastewaters", Unpublished paper.
3 -6
-------
Chapter 4
PHOSPHORUS REMOVAL BY MINERAL
ADDITION BEFORE THE PRIMARY SETTLER
4.1 General Considerations
Removal of phosphorus from wastewater can be accomplished by modifying the
conventional primary treatment process to include chemical precipitation with
aluminum or iron salts (1,2). The advantages of phosphorus removal during primary
treatment include flexibility in chemical feeding, adequate reaction times and mixing
conditions, flocculation and removal, of more suspended solids and suspended BOD in
the primary settler, and reduced loading of suspended solids and BOD to secondary
treatment processes. Significant increases in removals of both suspended solids and
BOD can occur concurrently with phosphorus removal. Phosphorus removal by
chemical addition is amenable to automatic control using process instrumentation to
measure waste flow and/or phosphorus concentration.
The disadvantage of addition of metal salts during primary treatment alone is that a
significant amount of the phosphorus may not be in the ortho form and may not be
easily precipitated. Higher phosphorus removals can be obtained with chemical addition
beyond primary treatment and with chemical additions at several points.
The available precipitation-flocculation processes (3,4,5,6,7,8) include addition of
ferrous or ferric chloride, ferrous or ferric sulfate, aluminum sulfate, or sodium
aluminate, followed by flocculation using an anionic polymer to enhance solids
separation. A strong base is often added between additions of ferrous iron and
polymer to counteract the depression of pH by the ferrous compounds.
In order to provide effective chemical additions to primary treatment units, the following
sequence is recommended:
1. Salt addition to raw sewage with thorough mixing. If a base is to be used,
it should be added at least 10 seconds later.
2. Reaction for at least five minutes.
3. Addition of anionic organic polymer followed by flash mixing for 20 to
60 seconds.
4. Mechanical or air flocculation for 1 to 5 minutes.
5. Gentle delivery of flocculated wastewater to sedimentation units.
Mixing may be carried out in specially designed basins or advantage may be taken of
existing high turbulence areas such as in a Parshall flume. Flocculation may similarly
be carried out in specially constructed equipment or may be allowed to take place within
some part of an existing system. Schematic diagrams of 2 plants, which demonstrate how
existing equipment can be used with mineral addition, are shown in Figures 4-1 and 4-2.
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The points of chemical addition will vary depending on the particular situation. It is
important, however, that addition points be downstream of return streams, such as
digester supernatant, if a high degree of phosphorus removal is required.
Optimum conditions for phosphorus removal processes are best determined in two steps:
a laboratory feasibility study followed preferably by full-scale plant trials if the plant
exists, or by pilot trials. The feasibility study should be a short-term laboratory evaluation
of various chemical systems to obtain preliminary estimates of chemicals required to
meet established phosphorus removal standards. During the plant trials, chemicals and
equipment should be supplied for full-scale operation over a period of 30 days or longer,
to obtain detailed design data.
Chemical removal of phosphorus from wastewaters during primary treatment is a
practical and controllable process which can result in low discharges of phosphorus
from existing primary treatment plants and secondary treatment plants of either the
trickling filter or activated sludge type. Additional costs for chemicals and attendant
mixing, dispersing, and monitoring equipment are moderate.
Data and correlations given in this chapter will provide guidance and permit making
design estimates. However, it would be risky to eliminate on-site experimental studies
because of variations in equipment design and operation, and in wastewater
characteristics.
4.2 Phosphorus Removal Data for Alum Addition
Alum has been used as a source of aluminum for phosphorus removal in raw
wastewater. Sodium aluminate can also be used if the natural alkalinity of the raw
wastewater is too low to prevent excessive reduction of pH by alum. However, no
investigations using aluminate have been reported. Tests have been conducted on
laboratory, pilot, and plant scales to determine the feasibility of alum addition to raw
wastewater.
Bench-scale tests of alum addition were conducted at Springfield, Ohio (9) and Two
Rivers, Wisconsin (10). All bench testing was conducted in the following manner.
Samples of wastewater were stirred on a six-place stirrer for approximately 2 minutes
at 100 rpm, followed by 18 minutes of gentle stirring (40 to 50 rpm). Alum additions
were made during the rapid mix period. When polymers were used, they were added
approximately 2 minutes after the alum addition. Treated samples were allowed to
settle for 30 minutes before total phosphorus analyses were made.
The respective states have required phosphorus removals of 80% at Springfield and 85%
at Two Rivers. At Springfield, an average A1:P weight ratio of 1.9:1 (molar ratio of
2.2:1) was required to reduce the average influent total phosphorus of 5 mg/1 by 80%.
Raw wastewater at Two Rivers required an average A1:P weight ratio of 0.93:1 (molar
ratio of 1.07:1) to reduce the average influent phosphorus of 20.1 mg/1 by 85%. These
test data indicate that as phosphorus concentration increases, the A1:P ratio necessary
4-4
-------
to obtain a specific per cent removal decreases. Individual tests at Two Rivers, for a
broad range of influent phosphorus concentrations, also showed the same variation.
Influent total phosphorus ranged from 7.6 to 28.5 mg/1, and the A1:P weight ratios
necessary to remove 85% ranged from 2.0:1 to 0.58:1, respectively.
A study of alum addition for phosphorus removal was conducted at Lebanon, Ohio on
both pilot and plant scales (11) to determine whether chemical treatment of raw
wastewater could be used successfully to precede activated carbon adsorption for
organic removal.
A 15 gpm pilot plant was utilized to determine the most effective dosages of alum to
be studied in the full-scale tests. Data from the pilot plant indicated that liquid alum
dosages of 150, 250, and 350 mg/1, as A^SO^-j- HF^O, and anionic polymer
dosages of 0.25 and 0.5 mg/1 should be evaluated. Only the Parshall flume,
preaeration basin, and primary clarifiers of the 1.5 mgd Lebanon plant were used in
this study. Alum was fed to the raw wastewater at the Parshall flume to assure
complete mixing. The preaeration basin was used for flocculation, with polymer being
added toward the end of the basin. The primary basins served as settlers for the
precipitated phosphorus. Figure 4-1 shows the chemical addition arrangement at
Lebanon which was used for both alum and iron studies. The iron experiments will be
covered later in this chapter.
Each dosage was evaluated for five days. A polymer dose of 0.5 mg/1 was found to
be satisfactory for improving floe formation and settling. An alum dose of 250 mg/1,
A1:P weight ratio of 3.4:1 (molar ratio of 3.9:1), removed 93% of the phosphorus in
this high alkalinity (400 mg/1 as CaCOo) wastewater. Effluent phosphorus
concentration did not exceed 0.7 mg/1 and averaged 0.5 mg/1. The alum dose of 150
mg/1, A1:P weight ratio of 1.5:1 (molar ratio of 1.7:1), removed only 74% of the
phosphorus; while the alum dose of 350 mg/1, A1:P weight ratio of 3.9:1 (molar ratio
of 4.5:1), reduced the phosphorus by 94%. The alum dose needed to meet the State
removal requirement of 80% would fall between the 150 and 250 mg/1 dosages
studied.
4.3 Phosphorus Removal Data for Iron Addition
Operating data from several plants utilizing mineral addition before the prynary clarifier for
phosphorus removal are shown in Table 4-1.
With primary treatment alone, phosphorus removals varied from 60 to 91%. In secondary
plants, iron addition before the primary resulted in phosphorus removals from 75 to 93%
over the entire plant. Iron doses varied between 10 and 90 mg/1, as Fe. The Lebanon study
showed that doubling the iron dose (from 45 to 90 mg/1) with the same polymer addition
raised phosphorus removal from 68 to 91%. Polymer doses in all of the iron studies ranged
from 0.26 to 0.7 mg/1.
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Iron carry-over occurred at Mentor, where FeC^ addition alone resulted in an effluent
Fe concentration of 42.5 mg/1. Lime addition reduced the iron leakage to an effluent
Fe concentration of 11.6 mg/1. Polymers were ineffective in reducing iron leakage.
Increased suspended solids and BOD removals were also obtained with iron additions.
At Grayling, Michigan, suspended solids removal was increased by 27% to an average
of 78%. At Benton Harbor, Michigan, where waste activated sludge is returned to the
primary, iron addition gave a 3.0% increase in solids removal over the entire treatment
plant. An example of increased BOD removal is seen in the Grayling, Michigan study.
Removal of BOD in the primary was increased by 45% after iron addition. This
resulted in a total BOD removal of 58% in the primary.
Typical values for primary settler detention times were 2.3 hours at Mentor and 2.25
•3
hours at Benton Harbor. At Mentor, air at a r.ate of 42.5 to 85 ft /minute was used
for flocculation in the mixing area of the primary tank.
Although lab test procedures should be adjusted to best approximate individual plant
conditions, outlines of procedures used at two locations may be useful as guidelines.
Jar tests were conducted at Mentor in the following manner. Iron, lime, and polymer
were added one minute apart to 1,000 ml of raw wastewater at a stirrer speed of 100
rpm. The chemicals and raw wastewater were then mixed for five minutes at 20 rpm,
and finally the mixture was allowed to settle quiescently for 1/2 hour before a portion
of the supernatant was pipetted off for analysis.
4.4 Overall Performance of Treatment Plants with Mineral
Addition to the Primary Settler
4.4.1 Primary Treatment Alone
The relative efficiencies of primary and secondary treatment with and without primary
phosphorus removal are summarized in Table 4-2. Phosphorus removal without
chemical precipitation in the primary clarifiers ranges from only 10 to 20% through a
conventional secondary treatment plant. At least one half or more of this removal is
due to sedimentation of the insoluble phosphorus fraction in the raw wastewater
during primary clarification. The addition of the precipitation-flocculation process in
primary clarification will result in 70 to 90% removal of total phosphorus in the
primary. These removal levels can be expected only in well designed primary clarifiers
that are not hydraulically overloaded. Suspended solids removals in the primaries will
be increased from 40 to 70% to 60 to 75%. This increase occurs despite the additional
solids generated by phosphorus removal. BOD removals in the primaries will be
increased from 25 to 40% to 40 to 50%. The additional removal is primarily due to
the suspended BOD fraction.
4.4.2 Effect on Secondary Treatment
Additional phosphorus removal through either trickling filter or activated sludge units
following chemical precipitation in the primaries can amount to 10 to 15%. This
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removal is attributed to biological flocculation. Split addition of metal in both the
primary and secondary portions of the plant should be considered if exceptionally high
levels (>90%) of total phosphorus removal are desired. Addition points for the
secondary metal feed at or near the effluent end of the biological treatment unit
should be included in the design.
Increased capture of suspended solids in the primaries results in reduced loading to the
secondary units. Although the solids entering the secondary are reduced, the liquid
volumes remain the same and, therefore, normal detention times should be used.
Essentially the same amount of soluble BOD must be insolubilized and removed. The
additional removal of suspended solids and BOD across the primaries will reduce the
amount of waste activated sludge to be handled. Savings in plant operating costs may
result from reduced air requirements for aeration, decreased chlorine demand, and
improved sludge filterability (28).
4.5 Dosage Selection and Correlation
The material presented in Chapters 6 and 7 on choice of chemicals and dosage in
trickling filter plants and activated sludge plants generally applies also to mineral
addition in primary plants. Therefore, the reader is referred to these chapters for
discussions of requirements for influent phosphorus data, procedures for obtaining
dosage design data, calculations of daily doses, and scheduling of dose rate changes to
match phosphorus variation. The reader is also referred to Chapter 10 for general
information on various chemicals. Material presented here will be limited to comments
and/or data on dosage selection, dosage correlation, and influent phosphorus variations.
It is generally desirable to plot or correlate test results on a simple basis to make
them more useful. A reasonably good correlation can be obtained by plotting the log
of the fraction of soluble phosphorus remaining as a function of the weight ratio of
metal to soluble phosphorus. Typical of such plots is Figure 4-3, where both the
fraction of soluble phosphorus remaining and the percent removed are plotted on the
log scale. The data in Figure 4-3 are from several ferric chloride phosphorus removal
studies conducted in the Great Lakes region.
Since there is usually considerable scatter in the data, such correlations should be used
very conservatively as a guide for predicting doses. On-site tests under actual plant
conditions are always preferred as the basis for such a plot. The required ratio of
metal ion to influent soluble phosphorus for a specified phosphorus removal can be
obtained from such a plot. This ratio is multiplied by the soluble influent phosphorus
concentration in mg/1 to determine the required chemical dose in mg/1.
As will be discussed in Chapter 6, the dose rate should be varied during the day to
match more closely the varying influent phosphorus concentration. The need for doing
this is illustrated by Figure 4-4 where some typical data on the diurnal variation of
ortho phosphorus in plant influent are presented.
4-10
-------
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AVERAGES FOR 5-DAY WEEK INFLUENT TO PLANT
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FIGURE 4-U DIURNAL VARIATION OF ORTHO PHOSPHORUS
4-12
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4.6 Case Histories
This section presents information on design and operation of three treatment plants which
remove phosphorus by adding metals before primary settling and by adding polymers as
required to supplement these metals.
The plant at Pontiac, Michigan employs commercial grade ferric chloride as a metal source;
Trenton, Michigan uses waste pickle liquor; Waukegan, Illinois has successfully used three
liquid forms of metals: ferric chloride, alum, and waste pickle liquor. Operations have been
effective at all three plants, and details are presented here to illustrate the variety of
approaches to phosphorus removal which are presently being applied successfully.
4.6.1 Pontiac, Michigan
The Auburn Plant in Pontiac, Michigan is a conventional activated sludge treatment plant
with an average design flow of 10 mgd. The facility was constructed in 1963 according to
design parameters presented in Table 4-3. Very simple chemical feed facilities were added to
permit addition of ferric chloride and polymers to the primary section of the plant as shown
in Figure 4-5.
A test program in 1969 (21) showed that feeding ferric chloride and polymer ahead of the
primary clarifier was an effective procedure to remove 80 to 90 percent of total phosphorus
in the wastewater. Long-term analysis of concentrations and forms of phosphorus in plant
influent indicate soluble compounds represent 60 to 80 percent of the total phosphorus
present, and that orthophosphates represent 35 to 45 percent of total phosphate. There
appeared to be minimal iron leakage in the effluent, and overall plant performance, in terms
of BOD and SS removals was improved by this technique.
The best point for addition of ferric chloride was just before the aerated grit chamber at the
head of the plant. Theoretical displacement through the grit chamber was approximately 10
minutes, but observations indicated actual detention was more on the order of 2 to 5
minutes at average flow rates. This was sufficient time, however, for mixing and initial
flocculation of the metal salt. Passage through the grit unit provided adequate lag time
before polymer addition.
Since 1972, the plant has been operating approximately as shown by the performance data
in Table 4-4 (29).
These data, mostly derived from 1973 experience, reflect a high-quality effluent and
well-stabilized sludge. When feeding chemicals, the primary section of the plant removes 40
to 60 percent of incoming BOD, 50 to 75 percent of incoming SS, and 50 to 75 percent of
incoming total phosphorus.
4-13
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Table 4-3
DESIGN DATA FOR PONTIAC, MICHIGAN
Description
Design Flow, mgd
Average Dry Weather 10
Peak Dry Weather 20
Maximum 36
Primary Clarifiers
Detention Time, 1 hr 1.5
Overflow Rate, 1 gpd/sq ft 1,000
Sludge Solids Concentration, percent 6.5
Sludge Volume, percent of total flow 0.08
Aeration Tanks
Detention Time,^ hr 6.5
MLSS, mg/1 2,000-3,000
Return MLSS, percent 40
Final Clarifiers
Detention Time,^ hr 1.9
Overflow Rate,1 gpd/sq ft 790
Sludge Digesters
Total Active Volume, cu ft 19 5,000
Vacuum Filters
Total Area, sq ft 375
Nominal Capacity, psf/hr 4
' Based on Average Flow
There has been no evidence of serious problems related to escape of colloidal iron from the
treatment system. Some 2 to 3 mg/1 of soluble or colloidal iron is found both in plant
influent and effluent.
Both ferric chloride and polymer are needed in this operation for good overall performance. If
iron dosage is reduced below the indicated mole ratio of 1.7, or if polymer addition is
discontinued, effluent quality begins to degrade in a matter of a few hours.
4-14
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Table 4-4
OPERATING DATA FOR PONTIAC, MICHIGAN
Description
Flow, mgd 10
Influent Characteristics
BOD,mg/l 100
SS,mg/l 140
Total P, mg/1 5
Chemical Addition
Iron Dose, mg/1 15
Iron Dose, Fe/P 1.7
Polymer Dose, mg/1 0.3
Overall Plant Performance
Effluent BOD, mg/1 10
Effluent SS, mg/1 20
Effluent Total P, mg/1 1
BOD Removal, percent 90
SS Removal, percent 86
Total P Removal, percent 80
Digesters
Temperature, deg F 90
Solids Loading, Ib VSS/cu ft/day 0.09
Digested Sludge Concentration, percent 7.2
Digested Sludge Volatile Fraction, percent 45
Vacuum Filters
Loading, psf/hr 3-4
Cake Solids Concentration, percent 30-35
Volume of sludge produced by chemical addition is approximately the same as generated
during conventional treatment. The chemical sludge is more concentrated but there have
been no problems with digestion or dewatering of sludges produced in this location.
Digested sludge is conditioned with ferric chloride and lime (at about the same dosage rates
as previously required for conventional sludge treatment) prior to vacuum filtration. Cake
forms readily on the filters but at modest yield rates. Both the vacuum filters and the
4-16
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multiple hearth furnace which follows are operated continuously for periods of 60 to 75 hr
during midweek. The units are shut down, cleaned, and held on standby the rest of the time.
Cost of treatment chemicals in 1973 was 2<^/ 1,000 gal of wastewater treated. Of this cost,
85 percent was for ferric chloride. Exact cost records are not available on capital equipment
installed in this system. That modest equipment outlay consisted of fiberglass tanks and
chemical metering pumps. Although quite workable, this temporary chemical feed
arrangement is to be replaced by permanent improvements in a 1 974 plant expansion.
4.6.2 Trenton, Michigan
Since December 1 970, Trenton, Michigan has operated a 5.5 mgd activated sludge treatment
plant with addition of pickle liquor and polymer for phosphorus removal and effluent
polishing (19,30). Design data for the Trenton plant are listed in Table 4-5.
Pickle liquor, the only metal salt used on a long-term basis at Trenton, is stored in two
rubber-lined steel tanks having a capacity of 25,000 gal each. These tanks have hot water
heating coils to keep coagulant temperature high enough to prevent crystallization of solids
during cold weather. The storage tanks are filled by tank trucks bringing pickle liquor from
a nearby steel plant; the trucks discharge into the storage tanks by applying low pressure
compressed air. The liquid metal salts are fed, by gravity, in proportion to the rate of
wastewater flow through the plant at a preselected concentration of metal. The
concentration can and is changed from time to time to meet changing needs.
Two polymer tanks, each holding 500 gal, are outfitted with variable flow metering pumps
which can be manually set to predetermined rates. These inject polymer between the
activated sludge aeration tankage and the final clarifiers. Mechanical mixing equipment is
provided in a small box between these two operations, and medium energy mixing is
provided there to blend polymers into the main flow.
Predesign pilot scale testing and plant scale startup operations indicated the system
performed best in the mode illustrated by Figure 4-6. Pickle liquor was added just ahead of
the aerated grit chamber, and the polymer was added to the activated sludge effluent en
route to the final clarifier.
Both the polymer addition and tube settlers installed in the final clarifier were intended to
improve collection and capture of fine floe observed in primary effluent during studies in
the late 1 960's.
Table 4-6 shows performance data typically observed at this treatment facility since it has
been in full scale service.
Subsequent to these operating data, an additional inflow of industrial waste amounting to 1
mgd has been added to the plant flow. The treatment system is presently accepting this
4-17
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Table 4-5
DESIGN DATA FOR TRENTON, MICHIGAN
Description
Design Flow, mgd
Average Dry Weather 5.5
Peak Dry Weather 6.5
Maximum Rate 16.5
Primary Clarifiers
Detention Time, hr 1.3
Overflow Rate, ^ gpd/sq ft 1,440
Weir Rate,l gpd/ft 23,000
Aeration Tanks
Detention Time, hr 5.5
Return Flow, percent of average 30
Final Clarifiers (Tube Module Type)
Detention Time, hr 3.5
Tube Overflow Rate, gpd/sq ft 1,200
Total Surface Overflow Rate, gpd/sq ft 500
Sludge Production
Primary (dry), Ib/day 4,000
Secondary (dry), Ib/day 5,000
Chemical (dry), Ib/day 4,000
Primary plus Chemical, percent 5
Secondary, percent 1.0-1.5
Gravity Thickening
Solids Loading, psf/day 6.2
Vacuum Filtration
Area, sq ft 360
Nominal Capacity, psf/hr 2-3
* Based on Average Flow
Including Return Flow
4-18
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additional waste, plus other increased domestic loads, so that average daily flow is now 7.5
mgd. This additional waste includes an increase in BOD loading to the plant of about 4,000
Ib/day. Plant performance at this increased load is reported as essentially the same as in
Table 4-6. Preliminary operating data from early 1974 indicate removal of 90 percent of the
BOD and SS, and 80 percent of total phosphorus. Sludge processing has continued on a
stable basis as indicated in Table 4-6.
Total cost of the 1974 facilities was $4.8 million. In this major construction project, it was
not possible to separate the relatively minor expenses related to the chemical treatment
facilities themselves. Something on the order of $5,000 could be identified for expenditures
Table 4-6
OPERATING DATA FOR TRENTON, MICHIGAN
Description
Flow, mgd 5.5
Influent Characteristics
BOD, mg/1 220
SS, mg/1 300
Total P, mg/1 9.5
OrthoP, mg/1 5.1
Chemical Addition
Iron Dose, mg/1 20
Iron Dose, Fe/P 1.2
Polymer Dose, mg/1 0.4
Overall Plant Performance
Effluent BOD, mg/1 10
Effluent SS, mg/1 10
Effluent Total P, mg/1 1
Effluent Ortho P, mg/1 0.5
BOD Removal, percent 95
SS Removal, percent 97
Total P Removal, percent 89
Ortho P Removal, percent 90
Gravity Thickeners
Thickened Sludge Concentration, percent 5.1
Thickened Sludge Mass, Ib/day 16,500
Cake Solids Concentration, percent 15-20
4-20
-------
for specific hardware, but concrete, electrical service, and other such features could not be
realistically identified as if they were an independent capital improvement. At any rate,
facility costs would be very low when capitalized over a period of years.
In 1973, the total operating cost of this facility was approximately 2Q
-------
4.8 References
1. Duren, J. W., "Detergent Builders and the Environment", Presented to the
Chemical Marketing Research Association, Chicago, Illinois (February 24, 1971).
2. Schuessler, R. G., "Phosphorus Removal, From Domestic Wastewater or From
Detergents?", Presented at the 57th Annual Meeting of the Chemical Specialties
Manufacturers Association, New York, N.Y. (December 15, 1970).
3. Leckie, J., and Stumm, W., "Phosphate Precipitation", in Water Quality
Improvement by Physical and Chemical Processes, Edited by Gloyna, E. F., and
Eckenfelder, W. W., Jr., University of Texas Press, Austin, Texas, p 237 (1970).
4. Daniels, S. L., "Phosphorus Removal from Wastewater by Chemical Precipitation
and Flocculation", Presented at the American Oil Chemists' Society, 1971 Short
Course, "Update on Detergents and Raw Materials", Lake Placid, New York (June
16, 1971).
5. Johnson, E. L., Beeghly, J. H., and Wukasch, R. F., "Phosphorus Removal with
Iron and Polyelectrolytes", Public Works, 100:11, p 66 (1969).
6. Schuessler, R. G., "Phosphorus Removal - A Controllable Process", Chem. Eng.
Prog. Symp. Series, 67:107, p 536 (1971).
7. Wukasch, R. F., "The Dow Process for Phosphorus Removal", Presented at the
FWPCA Phosphorus Removal Symposium, Chicago, Illinois (June 19, 1968).
8. Wukasch, R. F., "New Phosphate Removal Process", Wat. and Wastes Eng., 5:9, p
58 (1968).
9. Harriger, R. D., and Hoffman, F. L., "Phosphorus Removal Tests — Springfield,
Ohio for Black and Veatch Consulting Engineers", Technical Service Department,
Allied Chemical Corporation, Industrial Chemicals Division, Syracuse, New York
(1971).
10. Harriger, R. D., and Hoffman, F. L., "Phosphorus Removal Tests — Two Rivers,
Wisconsin for Black and Veatch Consulting Engineers", Technical Service
Department, Allied Chemical Corporation, Industrial Chemicals Division, Syracuse,
New York (1970).
11. Feige, W. A., "Full Scale Mineral Addition at Lebanon, Ohio", Inhouse Report,
Environmental Protection Agency, Water Quality Office, Advanced Waste
Treatment Research Laboratory, Cincinnati, Ohio (1971).
12. Hathaway, S. W., and Smith, J. E., Jr., "Fundamental Design Concepts for the
Lime Stabilization of Lebanon Raw Sludge", Environmental Protection Agency,
Water Quality Office, Advanced Waste Treatment Research Laboratory, Cincinnati,
Ohio, Unpublished paper (1971).
4-2:
-------
13. Green, O., Eyer, F., and Pierce, D., "Studies on Removal of Phosphates and
Related Removal of Suspended Matter and Biochemical Oxygen Demand at
Grayling, Michigan, March — September 1967", Division of Engineering, Michigan
Department of Health, Lansing, Michigan, and The Dow Chemical Company,
Midland, Michigan (1967).
14. Leary, R. D. and Ernest, L. A., "Municipal Utilization of an Industrial Waste for
Phosphorus Removal": Presented at the 32nd Porcelain Enamel Institute Technical
Forum at the University of Illinois (October 8, 1970).
15. , "Studies on Removal of Phosphates and Related Removal of
Suspended Matter and Biochemical Oxygen Demand at Lake Odessa, Michigan,
May-October, 1967," Wastewater Section, Division of Engineering, Michigan Depart-
ment of Public Health, Lansing, Michigan, and The Dow Chemical Company, Midland,
Michigan (1967).
16. Johnson, E. L., Beeghly, J. H., and Wukasch, R. F., "Phosphorus Removal with Iron
and Polyelectrolytes," Public Works, 100:11, p 66 (1969).
17. Boggia, C., and Herriman, G. L., "Pilot Plant Operation at Warren, Michigan,"
Proceedings of the 43rd Annual Conference of the Michigan Pollution Control
Association (1968).
18. Connell, C. H., "Phosphorus Removal and Disposal From Municipal Wastewater," U.S.
EPA Report No. 17010 DYE (February, 1971).
19. Hennessey, T. L., "Pilot Studies, Design and Operation of a Modified Activated Sludge
Plant Using Pickle Liquor and Tube Settlers," Annual Meeting WPCF, Atlanta, Georgia
(October, 1972).
20. Wilson, Thomas E., Greeley and Hansen Engineers, Chicago, Personal Communication
(1973).
21. Hennessey, John, Jelinski, Robert, Beeghly, Joel H., and Pawlak, Timothy J.,
"Phosphorus Removal at Pontiac, Michigan," The Dow Chemical Company, Midland,
Michigan.
22. Harriger, R. D., and Zuern, H. E., "Michigan City, Indiana, Wastewater Treatment
Plant, August 15-October 22, 1971, Chemical Precipitation of Phosphorus," Allied
Chemical Corporation.
23. Cole, Don and Tveite, Paul, "Sodium Aluminate for Removal of Phosphate and BOD,"
Public Works, p. 86 (October, 1972).
24. Robson, C. Michael, Nickerson, Gary L., and Morrison, R., "Chemical Addition at Two
Trickling Filter Plants to Enhance Treatment Efficiency," 45th Annual Conference,
WPCF, Atlanta, Georgia (October, 1972).
25. Berthume, Don, City of Essexville, Michigan, Personal Communication (1973).
4-23
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26. Gaughan, D. M. and Alvord, E. T., "Phosphorus Removal by Ferrous Iron and Lime,"
Final Report by Rand Development Corp. for the County of Lake, Ohio and EPA's
Water Quality Office on Federal Grant No. 172-01-68, Prepublication Copy (1971).
27. Progress Reports, Mentor, Lake County, Ohio, FWPCA Grant WRPD 172-01-68, May
15, 1968 through May 31, 1969 and September 1, 1969 through July 31, 1970.
28. Voshel, D., and Sak, J. G., "Effect of Primary Effluent Suspended Solids and BOD on
Activated Sludge Production," JWPCF, 40:5, Part 2, p R203 (1968).
29. Hennessey, J., City of Pontiac, MI, direct communication.
30. Hennessey, T. L., Consulting Engineer, Trenton, MI, direct communication.
4-24
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Chapter 5
PHOSPHORUS REMOVAL BY LIME ADDITION
BEFORE THE PRIMARY SETTLER
5.1 Introduction
Use of lime to treat raw wastewater dates back to the mid-1800's. Observers recognized that
calcium formed a precipitate both with hardness and with phosphates. Under one patented
technique both phosphates and lime (CaO) were added at the treatment plant to increase
the amount of precipitate formed and to aid removal of SS (1).
More recent experimental efforts involved lime treatment of raw wastewater to remove
phosphorus. These are typified by an investigation in New York, in which feeding 300 mg/1
of slaked lime to a typical domestic wastewater removed 94 percent of the phosphorus to an
effluent concentration of 0.6 mg/1, reduced SS by 95 percent to 4 mg/1, and reduced BOD
by 71 percent to an average of 101 mg/1 (2).
Other studies have suggested that lime addition to primary treatment is an efficient
technique (3, 4). These studies were at bench-, pilot- and plant-scale and total phosphorus
removals obtained were 80 percent or more with concurrent reduction in BOD of 60 to 70
percent. In this work, sufficient lime was added to bring primary effluent to a pH of 9.5 or
10. This pH level was tolerated without problems in the activated sludge process which
followed; carbon dioxide generated by biological activity served to correct pH to near
neutral levels. Recirculation of lime sludge was very beneficial. With sludge reseeding, as
much as 60 to 75 percent reduction in overall lime requirements were obtained compared to
tertiary lime treatment. Reaction time required to achieve settleable lime-phosphorus
compounds was reduced from 1.0 hr without recirculation, to 0.25 hr with reseeding.
A number of other studies, from bench tests through plant-scale trials, helped to establish
lime addition to primary treatment as a viable alternate for phosphorus removal operations
(5,6,7).
The low-lime system described in Section 5.2.1 can provide 80 percent phosphorus removal
on a consistent basis. By adding tertiary filtration and facilities for addition of metal
coagulants and polymer, the technique can reduce total effluent phosphorus to near 1.0
mg/1.
The high-lime system described in Section 5.2.2 offers an even higher level of efficiency.
With flocculant aids and filtration, it can consistently produce final phosphorus levels below
1.0 mg/1.
5-1
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5.2 System Layout and Operation
Two basic systems are considered in this section. One involves relatively low lime dosages to
primary influent, with biological treatment serving both to recarboriate the wastewater and
remove additional phosphorus for metabolic needs. The other approach entails high-lime
treatment in the primary section, followed by pH adjustment with supplemental carbon
dioxide and biological treatment. The two systems are discussed separately in this section,
after the comments below which apply to both approaches.
Sludge disposal during lime addition will be more complex than in a conventional biological
plant. Lime sludge may not be compatible with existing disposal processes. Total sludge
weight may be doubled or tripled over conventional operation. Recovery of lime by
recalcining may be less attractive in this process because treatment in primary units yields a
sludge which is high in concentration of organic and extraneous inorganic materials, and
relatively low in recoverable lime compounds.
Lime treatment appeals most where incoming hardness and alkalinity are low because less
lime is needed and less sludge is produced. However, turbid effluent may develop and
settling aids may be required to supplement bioflocculation to insure a clear final
effluent (8). Recirculation of lime sludge has proven beneficial to the basic reaction taking
place. It promotes a more complete reaction, reduces scaling, and lowers lime demand.
Removal of phosphorus by lime addition in primary treatment is accompanied by increased
BOD and SS removal at that point. This reduces organic load on secondary facilities and
could aid in establishing nitrification.
In evaluating lime treatment systems, the effect of reduced carbon loads on their activated
sludge unit should be considered. Since lower organic loadings to the secondary treatment
units result in less cellular synthesis, a reduction in sludge wasting may be required.
However, if the sludge is retained in the aeration system (in an attempt to keep MLSS up to
some prescribed level) sludge age may extend to the point that settleability suffers and
phosphorus escapes by cell lysis.
Also, lime precipitation in the primary clarifier will increase the capture of silts or other
similar inorganic solids. These weighting agents will no longer be carried into the activated
sludge unit for sorption onto mixed liquor solids, and this may result in a fluffier biological
floe.
5.2.1 Low-Lime Treatment
The basic appeal of low lime addition in primary treatment stems from two fundamentals:
the chemical law of mass action and the need for phosphorus in biological systems.
In the first of these, lime precipitates more phosphorus during initial stages of their mutual
reaction than it does when pH has been elevated and phosphorus concentrations are quite
5-2
-------
low. This effective performance by lime usually continues through the first 75 percent of
phosphorus reduction, and that occurs before pH 10 when lime demands for softening
reactions increase. Overall lime demands become much higher when the process is operated
to reduce phosphorus to very low levels.
Following initial phosphorus removal by lime, in the primary facilities, biological treatment
removes more phosphorus as bacteria utilize it to support cell synthesis. Generally,
phosphorus is not a limiting nutrient in domestic biological treatment systems as long as
primary effluent contains at least 0.2 mg/1 of soluble phosphorus. Microorganisms demand
phosphorus to synthesize cells in direct response to the carbon concentration of the
incoming waste; in municipal wastewater, carbon is the limiting element, not phosphorus.
The result is lime precipitation can proceed at economical levels, then biological treatment
can further reduce phosphorus residuals to an acceptable effluent concentration.
The basic low-lime treatment system is shown schematically in Figure 5-1. One advantage of
this process is the opportunity to use existing primary clarifiers for phosphorus removal.
Capital expenditure, therefore, could be relatively small. The process requires equipment for
feeding and flash mixing lime. Flocculation usually occurs in the inlet zone of the clarifier
but separate facilities may be preferred for this process. The elevated pH of primary effluent
is reduced by recarbonation due to carbon dioxide produced in biological metabolism. The
low-lime system may not be appropriate at trickling filter plants unless high effluent pH is
tolerable, since recarbonation would be minimal in trickling filter units.
5.2.2 High-Lime Treatment
For the purposes of this manual, high-lime treatment is defined as the addition of sufficient
lime in primary facilities to achieve pH 11. The basic system layout is generally as shown in
Figure 5-1, for low-lime treatment. Recarbonation with stored carbon dioxide may be
necessary and facilities for this would normally be provided. Recovery of lime by
recalcination, an alternate feature discussed later in this Chapter, would frequently be
included in high-lime systems.
High-lime treatment is applicable when effluent quality requirements include special
provisions such as softening (for reuse), low levels of soluble compounds of metals,
improved virus removal, or reliable and consistent reduction of phosphorus below 1.0 mg/1.
Some form of tertiary solids removal, such as filtration, would usually be employed in these
plants.
High-lime treatment would normally be combined with other biological, physical or
chemical processes to provide an overall system of advanced waste treatment. If biological
treatment is included, the primary clarifier phosphorus residual should be high enough to
meet metabolic requirements.
High-lime treatment may result in a net increase in plant effluent levels of total dissolved
solids and alkalinity. See Section 5.3.2.
5-3
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FIGURE 5-1
TYPICAL FACILITIES FOR LOW-LIME ADDITION IN PRIMARY TREATMENT
INFLUENT
LIME-
MIX
1
PRIMARY
CLARIFIER
DE WATER
AND
DISPOSAL
SLUDGE
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SLUDGE
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5-4
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5.3 Case Histories
Two case histories are included in this section. One involves low-lime addition in primary
treatment, while the other is a high-lime system. Both installations are further cited in
Section 5.5, and it is intended that information in this section will give better perspective to
those later comments on sludges involved.
5.3.1 Newmarket, Ontario—Low-Lime Process
Lime fed into the primary section of an existing 2.4 mgd conventional activated sludge
treatment plant in Newmarket was the first permanent installation for phosphorus removal
in Ontario, Canada (5,9). Phosphorus removal facilities have been operational in
Newmarket since March, 1971.
The original treatment facility (constructed in 1963) included screening and grit removal in
an aerated tank with 5 to 10 minutes detention. The original design data and subsequent
operating data are shown in Table 5-1. A schematic of the Newmarket treatment facility is
shown in Figure 5-2.
Table 5-1
DESIGN AND OPERATING DATA FOR NEWMARKET, ONTARIO
Operation Design Data Operating Data
Flow, mgd 2.4 1.9
Primary Clarifiers
Detention Time, hr 1.6 2.0
Overflow Rate, gpd/sq ft 1,330 1,050
Aeration Basins
Detention Time, hr 8 10
F/M, Ib BOD/day/lb MLSS 1.0 0.81
Final Clarifiers
Detention Time, hr 2.4 3.0
Overflow Rate, gpd/sq ft 1,000 800
Chlorine Contact
Detention Time, hr 0.42 0.53
^When lime addition was established, efficient primary clarification reduced
F/M to 0.2 Ib BOD/day/lb MLSS; concurrently, one aeration basin was shut
down so effective aeration time was reduced to 6.7 hr.
5-5
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Sludge treatment includes anaerobic digestion of combined primary and waste activated
sludges. The heated primary digester (40 ft. in diameter and 21.3 ft. deep) has a capacity of
2,000,000 gal and gas mixing is employed to distribute the daily solids loading of 0.1 Ib/cu
ft. A secondary digester, of similar size, completes the digestion process. Ultimate disposal is
by wet spreading on farmland.
Lime storage is in a 25 ft diameter by 17 ft deep tank (mostly underground) which holds 30
tons of hydrated lime made up to near a 30 percent slurry by weight. To put lime into
storage, bulk calcium hydroxide is discharged directly to the storage tank along with slurry
water. The new slurry, after thorough blending, is sampled and titrated to determine actual
lime concentration. This type of storage also allows use of waste carbide slurry from a local
acetylene plant. Continuous slurry agitation is provided by either of two centrifugal slurry
pumps mounted beside two positive displacement lime feed pumps. Slurry pumps require
double mechanical seals for this severe service. Lime feed lines require flushing connections.
Lime feed is based on preselected dose rates with automatic interlock to rate of wastewater
inflow. Metered slurry is discharged into a new flash mix box, where inflow is mechanically
agitated for about 8 minutes.
Lime operating data reported in Table 5-2 is based on a 200 mg/1 slaked lime dosage. Overall
phosphorus removal at this dosage was greater than 80 percent.
Performance was slightly less effective at a slaked lime dose of 175 mg/1. Also, waste carbide
lime, at rates equivalent to 150 mg/1 of slaked lime, were reasonably effective in removing
phosphorus. However, the acetylene waste proved highly abrasive to pumps and also formed
heavy encrustations enroute to the biological unit. Carbonate scale problems were fewer
using slaked lime, but occasional hosing down of primary launders was still necessary.
Encrustation did not occur in aeration basins.
Discharge of relatively alkaline (average pH 9.6 with peaks to pH 11.5) primary effluent to
the activated sludge system does not appear to distress the biological process.
Waste activated sludge is returned to the primary clarifiers and the resulting sludge mixture
is drawn to anaerobic digestion. During conventional operation about 5 gal total
sludge/1,000 gal plant flow was produced. Primary underflow was about 4 percent solids, of
which 65 percent were volatile.
During lime addition, waste activated sludge production is reduced 50 percent due to the
lower organic loads on secondary treatment. Additional sludge from lime reactions keep
primary underflow at 5 gal/1000 gal plant flow. In this mode, however, the solids content
averages 9 percent or more, of which 36 percent are volatile. The pH ranges from 8 to 10 in
this sludge. Concentrations up to 15 percent solids were obtained when operators thickened
in situ by allowing sludge depth to build to 2.5 ft on the clarifier bottom. No problems
developed with mechanical scrapers, and sludge lines remained unplugged.
5-7
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Table 5-2
OPERATING RESULTS FROM NEWMARKET, ONTARIO
Conventional Treatment
Lime Treatment
Percent Percent
Influent Primary Final Removal Influent Primary Final Removal
BOD
ss
Total Phosphorus
Orthophosphate
Hardness
Alkalinity
PH
249
341
14.2
8.1
225
386
7.8
172
126
12.5
8.3
229
377
7.8
21
17
8.7
7.6
235
227
7.7
92
95
39
6
-
_
-
227
321
10.3
6.7
239
391
8.0
84
80
2.7
1.2
240
374
9.6
9
7
2.0
1.8
111
247
8.0
96
98
81
73
-
-
-
Note: All data are mg/1 except pH;hardness and alkalinity are expressed as calcium carbonate.
Chemical-biological sludge is pumped to digestion, as in prior conventional operation. The
anaerobic system has adapted to this type sludge and complete digestion is obtained.
Operation over a period of months has indicated that digester stability was enhanced when
sludge was thickened by holding a sludge blanket nearly 2.0 ft deep in the primary clarifier.
This brought sludge pH down to 9.5 or less on a consistent basis, and thickened the mass to
11 percent solids. Digested sludge has a pH near 7.2. Digestion produces normal amounts of
methane gas, but free carbon dioxide is reduced by consumption in recarbonation reactions.
Digester supernatant contains less than 5 mg/1 of soluble phosphorus.
Capital costs, in 1971 dollars, for the lime treatment additions are shown in Table 5-3.
Operating costs are almost entirely limited to the 200 mg/1 lime consumption. In this
instance, there is no increase in the cost of sludge handling because the volume of digested
sludge has remained unchanged.
5.3.2 Central Contra Costa Sanitary District (CCCSD), California-High-Lime
Process
Recent plant-scale tests of lime addition before primary treatment have been conducted in
Contra Costa County, California (10,11,12). A complete advanced treatment test facility
has been in operation since November 1971. Full-scale units provide tests of all elements of
a 30 mgd water reclamation system currently under construction and scheduled for start-up
in 1976. Since the purpose of the CCCSD water reclamation system is to reclaim wastewater
for industrial reuse effluent requirements are quite strict and are listed in Table 5-4.
5-8
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Table 5-3
CONSTRUCTION COSTS AT NEWMARKET, ONTARIO
Structures
Mechanical Work
Electrical Work
Equipment
Engineering
Miscellaneous
$23,000
20,000
8,000
5,000
9,000
4,000
Total $69,000
Table 5-4
EFFLUENT REQUIREMENTS FOR CCCSD TREATMENT
FACILITY
Limit
Constituent (mg/1)
Total Phosphorus 1-0
BOD 10.0
Total Nitrogen 5.0
Hardness (as CaCO3) 300
Alkalinity (as CaCO3) 255
TDS (maximum increment above water supply) 375
In order to demonstrate effectiveness of the planned 30 mgd treatment system, investigators
converted existing treatment facilities to various plant-scale forms of advanced waste
treatment. The final layout of the overall test facility is shown in Figure 5-3. Table 5-5 gives
pertinent data.
After screening, raw sewage passed a chemical addition point where lime was added and
flash mixed with recycled primary sludge. Ferric chloride and polymer addition were also
possible at the same point. Wastewater then flowed to the preaeration basin. Heavy grit
5-9
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FIGURE 5-3
SYSTEM LAYOUT AT CONTRA COSTA COUNTY
RAW SEWAGE
FERRIC CHLORIDE
LIME
FLOCCULATION AND
PREAERATtON
I
PRIMARY
SEDIMENTATION TANK
CHEMICAL
PRIMARY
EFFLUENT i
CO;
OX IDAT ION -
NITRIFICATION TANK
SECONDARY
SEDIMENTATION TANK
N2
SLUDGE RECYCLE
SLUDGE TO
SOLIDS
PROCESSING
AIR
RETURN
SLUDGE
WASTE SLUDGE
+* TO
RAW SEWAGE
•METHANOL
DENITR IFICATION
TANK
I
•MIXING
AERATED
STABILIZATION TANK
I
FINAL
SEDIMENTATION TANK
RETURN
SLUDGE
I
WASTE SLUDGE
—+- TO
RAW SEWAGE
CHLORINE CONTACT
T
ADDITIONAL TREATMENT
FOR INDUSTRY
5-10
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Table 5-5
BASIC DATA FOR CCCSD
Operation
Preaeration and Grit Removal
Average Water Depth, ft 11
Detention Time, hr 0.53
Design Flow/Tank, mgd 2.5
Air Supplied, cu ft/gal 0.10
Primary Clarification
Average Water Depth, ft 10
Detention Time, hr 1.6
Overflow Rate, gpd/sq ft 1,200
Design Flow/Tank, mgd 2.5
particles such as sand and unreacted lime settled out there. The gentle air agitation following
lime addition promoted the growth of large floe particles which were readily settleable.
Long-term operation of this air mixing facility indicated it was effective, and this approach
eliminated problems associated with mechanical flocculators in raw wastewater.
Waste activated sludge was returned to the preaeration section; there it mixed and settled
with the chemical and primary sludges. The system had no separate facilities for treatment
of waste activated sludge.
In the preaeration basin it appeared the carbon dioxide content of agitation air might, by
the process of recarbonation, work against the lime reactions required in the preaeration
unit. However, if all carbon dioxide in the air were transferred into solution, less than 0.5
mg/1 dissolved carbon dioxide would be introduced. This was considered insignificant.
5-11
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However, carbon dioxide in the preaeration air did cause scaling inside the orifices of the
coarse bubble diffusers. Initially the tank had to be drained every week or two to clean
plugged orifices. Installation of high-pressure air blowoff facilities extended that period to
two months. In the 30 mgd plant, swing diffusers will be provided so tanks will not have to
be taken out of service for diffuser maintenance.
The lime-treated effluent was then settled in the rectangular primary tanks under operating
parameters described previously in Table 5-5.
During the course of the studies lime dosage was varied, and supplemental coagulants were
utilized to modify effluent qualities. Data for various modes of operation are shown in
Tables 5-6 and 5-7.
Operation at pH 11 was not as effective in phosphorus removal as operation at 11.5 (see
Table 5-5). The phosphorus remaining in the primary effluent at pH 11 was 2.3 mg/1
compared with less than 1 mg/1 at the higher dose.
Ferric chloride was tested as a supplemental coagulant to enhance capture of colloidal
phosphorus precipitates. At pH 11, a ferric chloride dose of 14 mg/1 produced primary
effluent phosphorus of less than 1 mg/1 (see Table 5-6). Even the relatively low lime dose for
pH 10.2 was possible if compensated by ferric chloride dosage of 24 mg/1. In this operation
primary effluent total phosphorus averaged 0.68 mg/1.
Removals of BOD, SS, and other parameters were comparable in each mode of operation.
However, increase in hardness across the primary varied considerably. Operation at pH 11
with ferric chloride yielded the lowest effluent hardness.
The data in Tables 5-6 and 5-7 represent results when primary clarifier underflow solids
were recycled to the chemical addition point. This was done to improve precipitation
reactions involving calcium, magnesium, and phosphorus. Total solids in the lime reactor
were maintained in the range of 1,000 to 2,500 mg/1. Later with solids recycle discontinued,
the major performance change was in effluent hardness. During operation at pH 11 with
ferric chloride dose of 10 to 16 mg/1, hardness across the primary increased an average of 32
mg/1. Under similar chemical feed rates, hardness increase across the primary was 6 mg/1
when flocculator solids were in the range of 900 to 1,500 mg/1 TSS. When flocculator solids
were increased to 1,700 to 3,900 mg/1, hardness increase across the primary was 16
mg/1 (12).
Considerable sludge thickening occurred in the primary clarifiers. Concentration of
underflow sludge varied with a number of factors, including the nature of the raw
wastewater. When a nearby cannery was in operation, underflow solids concentration did
not exceed 4.5 percent when operating at pH 11. This seemed due to anaerobic action and
gasification of peach and tomato solids in the raw sludge. Otherwise, 5 percent total solids
in the underflow was maintained. Occasionally underflow at total solids levels of 9 percent
5-12
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Table 5-6
LIME TREATMENT RESULTS AT CCCSD (12)
Flow: 1.30 mgd
pH: 11.5
Ca(OH)2 500 mg/1
Total Phosphorus
BOD
SS
TOC
Turbidity (JTU)
Calcium Hardness
(as CaCO3)
Magnesium Hardness
(as CaC03)
Influent Control Primary Chemical Primary
mg/1 mg/1 % Removal mg/1 % Removal
9.4 - - 0.96 90
190 103 46 50 74
199 57 71 41 79
107 59 45 37 65
35 - 16
76 - - 168
96 - - 40
Flow: 1.12 mgd
pH. 11.0
Ca(OH)2 400 mg/1
Influent Control Primary Chemical Primary
mg/1 mg/1 % Removal mg/1 7< Removal
9.2 - - 2.3 75
192 121 37 60 69
195 57 71 47 76
118 68 42 48 59
- 40 - 26
76 156
103 - 59
Table 5-7
LIME AND METAL TREATMENT RESULTS AT CCCSD (12)
Flow: 1.19 mgd
pH: 11.0
Ca(OH)2: 400 mg/1
FeCl3: 14 mg/1
Influent Control Primary Chemical Primary
mg/1 mg/1 % Removal mg/1 % Removal
Total Phosphorus 9.5 - - 0.85 91
BOD 210 109 48 53 75
SS 305 69 78 27 91
'IOC 130 72 45 37 72
Turbidity (JTU) - 45 - 14
Calcium Hardness
(asCaCO3) 76 - 139
Magnesium Hardness
(asCaCO3) 90 - - 32
Flow 1.20 mgd
pH 10.2
Ca(OH),: 289 mg/1
FeCl3. 24 mg/1
Influent Control Primary Chemical Primary
mg/1 mg/1 % Removal mg/1 % Removal
9.4 - - 0.68 93
178 106 40 59 66
235 59 75 31 87
117 68 42 43 63
41 - 15
72 - - 148
90 - - 70
5-13
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were obtained. At about 7 percent total solids, coning and bridging occurred in hoppers of
these particular tanks and septic sludge zones developed.
Contrary to experience with high-lime operation, underflow solids at pH 10.2 could not be
thickened to more than 4.5 percent. Use of an anionic polymer at 0.25 mg/1 allowed
thickening to 6 percent. An alternate anionic polymer did not significantly improve solids
underflow and neither polymer enhanced the effluent quality.
The pattern of deposition of solids in the primary clarifier was studied. Most of the sludge
was found in the first three-eighths of tank length. The heavier solids settled out rapidly in
this zone and the rest of the tank served to settle finally divided solids. Continuously
operating scaper flights kept the bottom relatively clean. The rectangular primary clarifier,
with preaerated grit removal, seemed suitable for the lime precipitation operation.
Primary clarifier effluent flowed to the activated sludge nitrification-oxidation tanks
directly with no continuous intervening recarbonation stage. External CC>2 was added to
mixed liquor only when needed by vaporization form liquid storage. Normally, CC>2 was
only necessary when lime addition operations were at pH 11.5 or higher. In-process CC>2
generation was sufficient for complete recarbonation with primary operation at pH 11.
Supplemental CCh was injected to the first bay of the activated sludge tanks to avoid scaling
in the transfer channel. All CC»2 sources contributed to lower pH to 7.5 in a short time.
CC>2 generated in the biological process from the oxidation of carbon provided most of the
recarbonation requirements in all operating modes. The amount of CC»2 generated from
oxidation was calculated from a total carbon balance around the oxidation-nitrification
stage and typically amounted to 55 mg/1 of CC>2 for recarbonation.
The amount of CC»2 generated from the hydrogen ions produced during nitrification was
calculated from alkalinity destruction, and appeared to amount to some 115 mg/1 under
typical conditions. The denitrification operation, occurring in a subsequent unit, was not
involved in recarbonation of primary effluent.
Daily supplemental CC>2 requirements averaged about 25 mg/1 after the process was
stabilized at pH 11.5. CC>2 demand varied diurnally according to the strength of the
wastewater.
The CCCSD treatment facility was equipped with a high-pH alarm to warn operators of the
need for exterior CO2- The 30 mgd water reclamation plant will use a fully automated pH
control system.
It was recognized that, in this particular plug-flow activated sludge system, the highest pH
levels would occur in the first part of the aeration tankage. With operation in the
conventional mode, primary influent and return activated sludge blended at the influent
end. Thus a large part of the oxygen demand also occurred in the first part of the tank.
5-14
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Enough oxygen to match the load in the head end of this tank would maximize CC>2
generation there. In the 30 mgd reclamation plant, DO probes in each pass will allow varying
the diffused air input according to load.
An unusual feature of this plant was application of two-stage wet classification of raw sludge
by centrifuges. The first-stage centrifuge received primary clarifier underflow (thickened in
situ) at a high rate. It captured two-thirds of the solids present: the resulting cake was 50
percent solids by weight and contained 85 percent of the calcium carbonate in the feed
sludge. Feed sludge was not conditioned with chemicals. This first-stage separation provided
sludge for recalcination. The centrate from the first centrifuge proceeded to a second
centrifuge designed for efficient capture of solids remaining in the feed. When treated with
polymer a drier cake was delivered. This cake which contained most of the organics and
phosphorus precipitated in the primary clarifier, was incinerated. Further information on
sludge handling is presented in Section 7.5.
5.4 Other Experience
This section contains information on lime treatment plants under construction, on full-scale
studies, and on related information which may be of benefit to designers of phosphorus
removal facilities. This is not intended to be a complete listing, but should be representative.
5.4.1 Rochester, N.Y.—Low-Lime Process
Bench scale treatability studies for phosphorus removal preceded design of a 100 mgd
treatment plant at Rochester, New York (13, 14). Activated sludge treatment alone
removed about 20 percent of the influent phosphorus (8 mg/1 as P). This agreed reasonably
well with phosphorus required for biological metabolism.
Jar tests of chemical treatment were performed. These tests examined alternated chemical
procedures to achieve required 80 percent phosphorus removal. Lime, and ferric chloride
with anionic polymers, were tried on raw wastewater. Either 100 mg/1 of lime as CaO (pH
9.5) or the combination of 40 mg/1 FeCl3 and 0.5 mg/1 anionic polymer reduced
phosphorus 70 percent. This removal plus biological removal raised overall reduction to 80
percent or more. BOD and SS reductions in jar tests were about 50 percent and 85 percent
respectively. In this case, lime was selected due to: 1) overall economy, and 2) ease of sludge
handling in the plant configuration to be described.
Existing treatment plant facilities at Rochester include primary clarifiers, Imhoff tanks,
sludge holding tanks, coil filters, and sludge incinerators. Proposed improvements include
modification primary clarifiers to add flocculation sections, addition of new primary
clarifiers incorporating flocculation sections, conversion of Imhoff tanks to complete-mix
activated sludge units, and addition of chlorination facilities. Plant startup is scheduled for
late 1975. Detention in flocculation sections of existing clarifiers will be 20 minutes. In the
new primary clarifiers, flocculation time will range up to 40 minutes. Existing and new
5-15
-------
primary clarifiers will have overflow rates of 1,000 gpd/sq ft and detention times of 2.3 and
2 hr, respectively, at average design flow. A schematic diagram of the expanded plant is
shown on Figure 5-4. Plant design is based on the criteria shown in Table 5-8.
Lime will be added just after flow passes the comminutors. The pH will be monitored ahead
of the flocculation basins and lime addition will be automatically regulated to maintain pH
9.5.
Anionic polymer will be fed just ahead of the flocculation basins when such addition is
necessary to improve settling. Provisions for returning solids from the primary clarifiers to
the flocculation basins are included.
Although lime treatment and biological treatment should remove 80 percent of incoming
phosphorus, there will be additional facilities for adding alum to aeration tanks. This
subsystem may be used alone or may supplement lime treatment for greater phosphorus
removal.
Complete-mix activated sludge is expected to disperse the primary effluent rapidly,
permitting recarbonation by bacterial CO2- F/M can be controlled over a broad range in the
configuration provided for aeration tanks.
Since 80 to 90 percent SS removal is expected in the primary clarifiers, the weight of
primary sludge (based only on increased SS removal) will be 1.5 to 2.0 times conventional
primary sludge production. Sludge produced by chemical precipitation must be added to
Table 5-8
DESIGN DATA FOR ROCHESTER, N. Y.
Average design flow, mgd 100
Maximum flow, mgd 200
Minimum flow, mgd 50
BOD loading, Ib/day 250,000
BOD concentration, mg/1 300
Suspended solids concentration, mg/1 300
Volatile suspended solids concentration, mg/1 240
Total hardness, mg/1 200
Phosphorus, mg/1 as P 8
5-16
-------
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5-17
-------
this, raising the total to about 3.0 times the normal amount. Primary sludge quantities at
design flow with lime treatment are estimated at 175 tons/day. Since organic load on the
activated sludge section will be reduced about 50 percent, activated sludge will be reduced
substantially. At design loads with lime treatment, waste activated sludge is estimated at 30
tons/day, compared to about 75 tons/day if lime treatment were not used. Solids handling
facilities have been designed for a capacity of 270 tons/day.
Sludges will be combined and gravity thickened to a concentration of 5 percent. Thickeners
are designed for a solids loading of 13 Ib/day/sq ft. After thickening, sludge will be
dewatered on vacuum filters and incinerated in multiple hearth furnaces. Incinerator ash will
be sluiced and pumped to two 0.5-acre lagoons, each having a water depth of 11 ft,
operating in series. Effluent from lagoons will be pumped back to the head of the plant.
Lagoons will probably have to be emptied every 9 to 12 months.
5.4.2 Lower Allen Township, Pa.-Low-Lime Process
Recalcination of sludge generated by lime addition in the primary is one feature of a new
6.0 mgd plant at Lower Allen Township, Pennsylvania (15). This plant, where startup began
in February 1974, is a complete-mix activated sludge system with low-lime addition in the
primary section.
Rough screening, degritting, and comminution precede lime injection. Flash mixing blends
both recalcined lime and recycled seed sludge form the primary clarifier. Lime addition is
controlled automatically within the range of pH 9.4 to pH 9.8. Mixed flow is divided
between two flocculating clarifiers, each 85 ft in diameter and 12 ft in depth. These have a
design average overflow rate of 700 gpd/sq ft.
Primary effluent is treated biologically in four activated sludge basins, each 45 ft, x 90 ft x
15 ft. Recarbonation reduces pH rapidly to between 7.5 and 8.5. Biological uptake is
expected to complement lime precipitation in order that overall phosphorus reduction will
be 80 percent. Phosphorus concentration in influent wastewater is 9.0 mg/1. Two final
clarifiers are each 85 ft in diameter and 12 ft in depth.
Combined primary and waste activated sludges are blended and gravity thickened. After
storage, sludge is conditioned before being fed to two parallel solid-bowl centrifuges.
Dewatered cake from the centrifuges is fed to a multi-hearth dual purpose furnace. In the
furnace, organic fractions are volatized while waste inert inorganic sludge is dried. The
furnace is also designed to recalcine and recover quicklime for reuse at the head of the plant.
5.4.3 Manhattan, Kansas-Low-Lime Process
In Manhattan, Kansas, research into phosphorus removal using low lime and biological
treatment began with laboratory scale investigations (16). It was found that both
orthophosphates and polyphosphates could be precipitated readily. Complex phosphates
5-18
-------
present in wastewater seemed to inhibit calcium carbonate precipitation, to the extent that
recalcination for lime recovery seemed impracticable. In these studies, 150 mg/1 of slaked
lime raised pH to 9.5. Total phosphorus removal before biological treatment was 60 percent,
and SS were reduced 90 percent. Lime-primary sludge production was about twice that
obtained in conventional operation and overall sludge production, including waste activated,
was about 1.5 times that found with conventional secondary treatment.
The low lime process was operated at a constant pH. Performance of the downstream
complete-mix activated sludge system was not disturbed. Microbial CC>2 production in the
activated sludge unit was sufficient to maintain pH near neutral. A mixture of
lime-precipitated primary sludge and waste activated sludge was dewatered and vacuum
filtered with small additions of anionic polymers.
A pilot scale demonstration of the technique followed (17). The wastewater was typical
domestic wastewater generated at the rate of 80 gal/capita/day. Untreated wastewater had a
pH near 7, and other characteristics were BOD and SS near 200 mg/1, total phosphorus of 10
mg/1, alkalinity of 235 mg/1, hardness of 150 mg/1, calcium of 110 mg/1, and magnesium of
40 mg/1. The pilot facilities included a 15,000 gpd conventional treatment process, with
lime addition to the primary. Influent wastewater was supplied to the system on a diurnal
pattern from a variable speed pump and lime slurry was injected into the discharge line of
that pump. Flow then entered a mixing and reaction tank where 'detention was near 2
minutes. Lime feed was made proportional to flow initially, but was later converted to pH
control.
Primary clarification followed with a surface loading rate of 225 gpd/sq ft. Primary sludge
was extracted hourly. Primary effluent flowed into plug flow aeration tankage where it was
detained about 5 hr. The final clarifier had an overflow rate similar to the primary clarifier
but sludge withdrawal was continuous.
Primary and secondary sludges were detained separately in holding tanks large enough to
store the sludge from an entire day's operation. This permitted more effective sampling than
could be obtained by compositing smaller portions. The two sludges were then blended into
a third tank for dewatering studies. The biological system operated to give sludge age of 12
to 15 days when MLSS was 2,000 mg/1.
During the pilot study, a number of the operational features were changed deliberately or
through circumstance. It appeared the most important operational variable was pH
following addition of lime to the primary effluent. If that pH was held between 9.5 and 10,
phosphorus removals of 75 percent or more were generally obtained in the primary unit (see
Figure 5-5). Consistent phosphorus reduction was not obtained unless the pH was kept at
9.8. When pH was increased to 11.5, 90 percent phosphorus removal resulted, but a
considerably greater lime dosage was required.
5-19
-------
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5-20
-------
It appeared that fine floe escaping the primary clarifier contained, in particulate form, most
of the 25 percent total phosphorus residual which entered the biological section. Reduction
of the flow rate (which was already quite low in terms of surface loading of the primary
clarifier) did not materially improve capture of precipitated phosphorus. In this instance a
higher degree of performance was not obtained by the use of polymer.
Phosphorus removal in the biological section was 5 percent or less of the total phosphorus
coming into the treatment plant.
Lime treatment to pH 9.5 to 10 resulted in primary BOD reductions near 60 percent, a
considerable improvement over the usual 30 or 40 percent obtained in plain clarification. SS
concentrations were not materially reduced, as already mentioned in connection with lime
floe which was escaping the primary unit. CC>2 production by bacteria in the activated
sludge unit neutralized final effluent to levels near pH 8.
Addition of lime to the primary clarifier resulted in physical and biological changes in the
activated sludge. Mixed liquor turned gold or light brown instead of the usual chocolate
brown characteristic of this operation. Bacteria in the culture tended to shift into
filamentous species. Among the scavenger organisms, population of free swimming ciliates
declined, while rotifers increased to dominance.
The technique produced two to three times more total weight of sludge than generated
during conventional treatment. However, with properly administered polymer, the sludge
could be concentrated. The most attractive sludge treatment systems involved polymer
conditioning, dewatering by vacuum filter or centrifuge, and incineration or land fill.
Primary sludge had a volatile fraction of 50 percent during lime addition, as opposed to 65
percent in conventional systems. Chemical sludge would compact to a 3 percent density in
0.5 hr. This concentration could be considerably improved with an anionic polymer.
Biological solids were produced at the rate of about 50 mg/1 as the flow passed through the
aeration tank. Biological sludge represented 7 percent by weight of the total sludge
produced when using the lime technique.
Total sludge resulting from lime treatment was estimated at 3,750 Ib/million gal of flow.
This compared with 1,500 Ib/million gal, of which some one-third would be waste activated
sludge. Sludge conditioning with anionic polymer (2.5 Ib/ton dry solids) yielded a filtration
rate of 5 Ib/sq ft/hr.
Several problems developed during the course of the study, and these should be considered
by designers of such plants. Milk of lime solution had a marked tendency to scale everything
it contacted prior to injection into the wastewater. Following injection, lime depositing out
of treated wastewater caused scaling problems on the membrane of a pH probe, and on
wetted parts of an automatic phosphorus analyzer. These and other surfaces also tended to
collect a biological slime. Lime scale on pH probes was removed every other day by
submersing them into diluted acid.
5-21
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This pilot plant used 60 deg. hoppers for sludge collection in the primary clarifier. They did
not work well with lime sludge which had a tendency to cling even to these sharply sloped
bottoms. In due time denitrification occurred within that trapped layer and floating solids
resulted. Observers noted that sludge also adhered to hopper walls in the final clarifier and
became anaerobic prior to withdrawal. It appeared that this resulted in release of cellular
phosphorus into the effluent.
5.4.4 Richmond Hill, Ontario—Low-Lime Process
Following laboratory and pilot scale studies, a three-month plant scale investigation of lime
treatment for phosphorus removal was conducted using a 0.2 mgd section of a 2 mgd
conventional activated sludge plant in Richmond Hill, Ontario (18). Lime was added before
the primary clarifier on a continuous basis and no special mixing or flocculating equipment
was provided.
Raw wastewater had alkalinity of 381 mg/1 and total hardness of 401 mg/1. Average total
influent phosphorus was 10.6 mg/1 (as P). Normal plant operation reduced total phosphorus
by 39 percent. BOD and SS removals in the primary clarifier were 21 and 37 percent,
respectively.
Lime dosage of 175 mg/1 produced a pH of 9.3. During lime addition, BOD and SS removals
in the primary clarifier averaged 72 and 78 percent, respectively. Total phosphorus removal
averaged 83 percent in primary treatment, and removal through the entire plant was 92
percent, resulting in a total phosphorus concentration in the secondary effluent of 0.9 mg/1.
Throughout the study, soft calcium carbonate scale accumulated on effluent weirs and
troughs of the primary clarifier but was easily sluiced away using a water hose. There was no
appreciable buildup on submerged surfaces within the primary clarifier.
Lime treatment did not upset the biological process in the test section. The pH of 9.3 in
primary effluent was promptly reduced below 8 by CO2 generated in the aeration tank.
MLSS equilibrium was reached and maintained throughout the study. Due to improved
primary performance, organic loading to the conventional (activated sludge system) was
reduced and its behavior shifted to operation as an extended aeration process with no
secondary sludge wastage.
Sludge precipitated in the primary clarifier was a mixture of calcium phosphate, calcium
hydroxide, calcium carbonate, and coagulated organics and grit. This sludge was free flowing
and considerably different from the gelatinous matter previously produced in bench scale
studies. Sludge removed from the primary unit had a solids content of 6 to 8 percent. It
gravity thickened readily to 10 to 12 percent solids, but was difficult to dewater without
proper chemical conditioning. Biological activity in the lime-sludge was inhibited by the
high lime content. The sludge proved free from noxious odors even after prolonged holding
under anaerobic conditions. Supernatant from the chemical sludge, after 5 weeks of storage,
5-22
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contained no appreciable amount of phosphorus, either total or soluble. It was concluded
that lime treatment could be readily controlled either in response to flow or on the basis of
pH in a flash mixer or primary effluent.
5.4.5 Seneca Falls, New York
A combined biological-chemical precipitation process (proprietary name "PhoStrip") for
phosphorus removal was applied in a short-term (30 day) study at the existing Seneca Falls,
N.Y. treatment plant in 1973. This process involves modification of the conventional
activated sludge process. A schematic of the Seneca Falls treatment scheme is shown in
Figure 5-6.
The major modification to the activated sludge process is the incorporation of the
phosphate "stripping" tank in the return sludge line. This tank permits the return sludge to
become anaerobic which, apparently, induces the sludge organisms to secrete phosphate.
The phosphorus enriched supernatant from the stripper is then dosed with lime and
returned to the primary clarifier for sedimentation of phosphorus. Data obtained during the
30-day trial are included in Table 5-9.
Reported chemical costs for the Seneca Falls trial, based on a lime dosage of 24.1 mg/1 (as
CaO), were $2.01 per million gallons of raw wastewater. Additional costs would include the
phosphate stripper tank, the lime feeding, storage and mixing facilities, and associated
pumping facilities. At Seneca Falls, a spare clarifier was used as the phosphate stripper tank.
It is claimed that total costs for the project due to chemical cost savings may result in sig-
nificant savings compared to conventional mineral addition.
5.5 Sludge
Detailed information in Chapter 11 concerns amounts of sludge generated during lime
addition, and offers design parameters required to select and size sludge treatment and
disposal facilities. Additional information specifically related to sludge produced by lime
addition in primary treatment operations is discussed in this section.
Plant scale centrifuge trials on underflow from lime treated primary units at Newmarket,
Ontario, were considered successful (19). Centrifugation captured some 99 percent SS in the
feed, and cake solids were near 30 percent by dry weight. The performance of the solid
bowl centrifuge was such that centrate contained 700 mg/1 SS and about 20 mg/1 total
phosphorus. Centrate was returned to the flash mixer preceding the primary clarifier.
Polymer requirements were less than 1 Ib/dry ton of solids. Overdosing of polymer caused
foaming in the centrate.
Sludge production during low lime addition at Manhattan, Kansas (17) was 250 percent of
the weight of sludge produced by conventional treatment. However, being primarily a grainy
chemical sludge, this mixture of primary and waste activated sludge was much easier to
handle, dewater, and dispose of than sludge produced during normal plant operation.
5-23
-------
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IT) CO
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5-24
-------
Table 5-9
RESULTS OF PHOSPHORUS REMOVAL TRIAL AT SENECA FALLS, N. Y.
Plant Flow, mgd
Design 1.75
Average 0.9
Influent
BOD, mg/1 158
P, mg/1 6.3
Detention Times (based on average flow), hours
Primary Clarifier t 3.9
Aerator 5.5
Secondary Clarifier 3.2
Phosphate Stripper 18
Return Flows, percent of raw flow
Sludge to Phosphate Stripper 24
Sludge to Aeration Tank 10
Supernatant to Primary Clarifier 14
Lime Dose, mg/1 as CaO 24.1
SS, mg/1
Mixed Liquor 1,440
Return Sludge to Stripper 7,840
Return Sludge to Aerator 15,910
Effluent
BOD, mg/1 3.5
P, mg/1 0.55
Plant Performance
BOD Removal, percent 98
P Removal, percent 91
5-25
-------
Canadian studies (5), taken as a group, indicate total sludge production at a conventional
plant will increase from 1800 Ib/million gallons to 4600 Ib/million gallons with addition of
some 200 mg/1 lime to primary clarifiers. Solids concentration increased from 4.5 to 9
percent, so the volumetric increase was 25 percent.
In a primary plant, sludge production increased from 800 Ib/million gallons to 2400
Ib/million gallons with solids concentration rising from 7 to 15 percent (a net 50 percent
increase in sludge volume).
Plant-scale testing of two-stage centrifugation of combined lime-treated primary and waste
activated sludge has been applied at Central Costa County Sanitary District (CCCSD),
California (11). This permitted a degree of sludge classification vital to the integrated system
of lime recovery and sludge disposal to be utilized in the 30 mgd treatment plant under
construction at CCCSD. Figure 5-7 depicts the solids handling and disposal facilities which
will be incorporated in the 30 mgd plant; Figure 5-8 depicts the plant-scale sludge handling
system used at CCCSD to develop design data for the full-scale system.
The purpose of the first of the two centrifuges is to maximize recovery of calcium carbonate
while rejecting other sludge components, such as organic material, phosphorus and
magnesium compounds, and other constituents coprecipitated with calcium carbonate. The
first-stage cake represents phase, all of which have a slower settling rate than calcium
carbonate, are deliberately carried out with the process centrate and passed on to a
second-stage centrifuge operated to obtain high solids recoveries. Second-stage solids are
incinerated (breaking down organic compounds in the process) and the total ash, including
fine inert materials, can be disposed of without recycling back into the wastewater
treatment process.
Optimum first-stage performance in the test system was obtained after varying critical
centrifuge variables: feed rate, centrifugal force, pool depth, and conveyor speed.
Observations were made on individual sludge constituent recovery throughout a broad range
of these process variables. Figure 5-9 shows the generalized results. The definition of
performance allowed setting the centrifuge variables to capture a high percentage of calcium
carbonate but a low percentage of all other sludge constituents.
Certain trends were noted in the CCCSD studies which may have application in other
locations. In general, results followed past practice in centrifugation of calcium carbonate
and, almost with exception, followed basic centrifuge fundamentals.
When wastewater was lime treated to a pH of 11 or 11.5, the highest percentage of calcium
carbonate was recovered in the first unit and the greatest reject of extraneous solids
occurred. Sludge produced when flocculating at pH 10.2 was not as easily classified.
Separation was most effective when pool depth in the bowl was at very shallow settings.
Dewatering of the calcium carbonate cake was materially improved by raising centrifugal
force to the range of 2,100 g's. Polymer preconditioning of sludge was not required.
5-26
-------
FIGURE 5-7 PROPOSED SLUDGE FACILITY AT CONTRA COSTA COUNTY
FROM
RECALCINING FURNACE
1 ' 1 '
MAKEUP
LIME
WASTE
B IOLOGICAL
SOL IDS
LIME
STORAGE
CENTRATE RETURN
PRIMARY
EFFLUENT
PR IMARY C L AR IF IER
FIRST
STAGE
T H ICKENER
WET
CL ASSIF ICAT ION
SECOND
STAGE
THICKENER
MULTIPLE
HEARTH
RF. CALCINE
FURNACE
GAS
SCRUBBER
RECALCINED
ASH
REJECTS
•2nd STAGE CENTRIFUGE
OR
FILTER PRESS
OR
VACUUM FILTER
CENTRATE TO
PRIMARY
QRY
CLASS IFICATION
RECYCLE LIME TO STORAGE
MULTIPLE
HEARTH
FURNACE
HIGH LIMC
ACCEPTS
ASH TO DISPOSAL
5-27
-------
FIGURE 5-8 PLANT SCALE SOLIDS TESTING SYSTEM AT CONTRA COSTA
RAW
PRIMARY
-*•
EFFLUENT
Centrote Processing Alternatives
PRIMARY CLARIFIER
Thickening
CENTRATE
T HICKENER
UNDERFLOW
SOLIDS
Centrifugation or
Classification
SLUDGE
TO
INCINERATOR
HIGH
CaC03
ACCEPTS
OR
WHOLE
SLUDGE
SLUDGE TO
INCINERATOR
GAS
SCRUBBER
2ND STAGE
CENTRATE TO
PRIMARY
CLARIFIER
MULTIPLE
HEARTH
INCINERATOR
Incineration or
Recalcination
RECALCINED PRODUCT
OR ASH
First-stage centrate passed to the second centrifuge for further treatment. Several polymers
were tried and some proved effective in increasing solids capture when fed at the rate of 2
Ib/ton of solids. Solids recovery of 80 percent and resulting cake dryness of 15 percent
solids were typical operating levels obtained after experimentation. Setting the machine for
maximum pool depth and operating near 2,100 g's gave satisfactory results in this particular
installation.
5-28
-------
FIGURE 5-9 SLUDGE CLASSIFICATION CURVES AT CONTRA COSTA
IOO
20
RECOVERY
OF
4O
TOTAL
60 8O
SOLIDS, PERCENT
too
The extensive investigations at this installation (11) have only been briefly summarized here.
The conclusion is that, in this instance, two-stage classification by centrifuges was
applicable. Results of the operation were generally predictable by centrifuge laws. Other
locations considering this type of operation should make comparable predesign studies.
It is also worthwhile to note that sludge production at the CCCSD test installation was
predicted by theoretical procedures similar to those described in Chapter 9. Actual sludge
production was then measured under four different operating modes, and operating
experience agreed within about 10 percent with predicted values. These predictions and the
actual production rates both included biological sludge from the nitrification-denitrification
system, as well as settleable solids normally removed in primary treatment. In this case, the
organic fraction of total sludge was one-quarter to one-third of the combined weight
involved.
5-29
-------
5.6 Costs
The capital cost of adding rapid mix and flocculation facilities to existing plants has been
estimated for 1,10, and 100 mgd plants. These should serve as a general guideline where
phosphorus is to be removed by lime addition to the primary clarifier. Estimated costs for
rapid mix facilities include concrete basins (with detention of one minute at design flow),
and mixers with shafts and impellers of stainless steel. Flocculation facilities consist of
baffle walls and mixers to provide a flocculation chamber with 30 minute detention within
existing circular clarifiers. Estimated costs of these are given in Table 5-10 and include
allowance for Contractor's installation, overhead, and profit, plus 20 percent for engineering
and contingencies. Figures given relate to an EPA-STP Cost Index of 185.
Table 5-10
TYPICAL CAPITAL COSTS FOR LIME TREATMENT IN PRIMARY
FACILITIES
1 mgd 10 mgd 100 mgd
Rapid Mix
Flocculation
$ 2,200
64,000
$ 13,000
264,000
$ 82,000
983,000
Total $66,200 $277,000 $1,065,000
Investment figures for lime storage and feed facilities are given in Chapter 10.
The cost/1000 gal for adding single-stage lime treatment to an existing plant was determined
by amortizing capital cost at 6 percent over 25 years and adding the cost of chemicals.
Operational expenses such as labor and power are not included.
Lime costs/1000 gal are based on a chemical cost of $22/ton with 90 percent available CaO
at a dosage of 150 mg/1 as CaO. The equipment and chemical costs (^/lOOO gal) are given in
Table 5-11.
5-30
-------
Table 5-11
TYPICAL TOTAL COSTS FOR LIME TREATMENT IN PRIMARY
FACILITIES
Equipment Amortization
Rapid mix and flocculator
Chemical storage and feed
system
Chemical Cost
Total Cost (Hi, 000 gal)
1 mgd
1.4
0.4
1.5
3.3
10 mgd
0.6
0.2
1.5
2.3
100 mgd
0.2
0.1
1.5
1.8
Costs do not include allowance for equipment for sludge handling and disposal and
corresponding operating costs. Since a considerable quantity of additional sludge is
produced, this expense is likely to be substantial.
5.7 References
1. Pearse, et ah, "Chemical Treatment of Sewage," Sewage Works Journal, 7:6, p. 997
(1935).
2. Karanik, J. M., and Nemerow, N. L.. "Removal of Algal Nutrients," Water and Sewage
Works
3. Albertson, O. W., and Sherwood, R. J., "Removing Phosphates from Wastewater,"
Industrial Water Engineering, p. 30 (November, 1967).
4. Albertson, O. E., and Sherwood, R. J., "Phosphate Extraction Process," JWPCF, 41:8,
p. 1467 (1969).
5. Boyko, B. I., and Rupke, J. W. G., "Technical Implementation of Ontario's Phosphorus
Removal Programme," Proc., 28th Purdue Industrial Waste Conference, Lafayette,
Indiana (May, 1973).
5-31
-------
6. Spohr, Guenter, and Talts, Andres, "Phosphate Removal by pH Controlled Lime
Dosage," Public Works, p. 63 (July, 1970).
7. Tofflemire, T. J., and Hetling, L. J., "Treatment of a Combined Wastewater by the
Low-Lime Process," Journal Water Poll. Control Fed., 45, p. 210 (1973).
8. Bishop, Dolloff F., O'Farrell, Thomas P., and Stamberg, John B., "Physical-Chemical
Treatment of Municipal Wastewater," Journal WPCF, 44:3, p. 361 (March, 1972).
9. Black, S. A., "Lime Treatment for Phosphorus Removal at the Newmarket/East
Gwillimbury WPCP," Research Paper No. 2032, Research Branch, Ministry of the
Environment, Ontario (May 1972).
10. Parker, D., Brown & Caldwell Engineers, Private Communication (June, 1973).
11. Parker, D. S., Zadick, F. J., and Train, K. E., "Sludge Processing for Combined
Physical-Chemical-Biological Sludges," Grant #R801445, Office of Research and
Monitoring, EPA (March, 1973).
12. Horstkotte, G. A., et al., "Full-Scale Testing of a Water Reclamation System," Jour
WPCF, 46:1, p. 181-197 (January, 1974).
13. Wahbeh, V. N., "Design and Modification of the Rochester, New York Wastewater
Treatment Plant for Phosphorus Removal," Presented at the Technical Seminar/
Workshop on Advanced Waste Treatment, Chapel Hill, North Carolina (February,
1971).
14. Weller, L., Black and Veatch, Kansas City, Mo., direct communication (1974).
15. Paul, P., Gannett, Fleming, Corddry and Carpenter, Inc., Harrisburg, Pa., Direct
Communication (1973).
16. Schmid, L. A., and McKinney, R. E., "Phosphate Removal by a Lime-Biological
Scheme," JWPCF, 41:7, p. 1259 (1969).
17. Schmid, Lawrence L., "Pilot Plant Demonstration of a Lime-Biological Treatment
Phosphorus Removal Method," Water Quality Office, Environmental Protection
Agency, Project 17050 DCC (September, 1972).
18. and Lewandowski, W., "Phosphorus Removal by Lime Addition
to a Conventional Activated Sludge Plant," Div. of Res. Publ. No. 36, Ontario Water
Resources Commission (November, 1969).
19. Smith, A. G., "Centrifuge Dewatering of Lime Treated Sewage Sludge," Research Paper
No. W2030, Res. Br., Ministry of the Environment, Ontario (May, 1972).
5-32
-------
Chapter 6
PHOSPHORUS REMOVAL IN TRICKLING FILTERS
BY MINERAL ADDITION
A logical approach to removal of phosphorus in trickling filter plants would be the use
of precipitating chemicals somewhere in the system. Iron and aluminum can both be
used for this application. Because of their similar action, iron and aluminum salts are
commonly referred to collectively as minerals in wastewater treatment and are
considered separately from calcium (lime). Required physical facilities for adding iron
and aluminum are relatively dimple. They are covered, along with special considerations
for specific chemicals, elsewhere in this manual. This discussion will emphasize the
merits of alternative addition points, review methods for determining dosage, and report
studies where chemicals have been dosed directly to a trickling filter and to trickling
filter effluent. Broader aspects of mineral addition for phosphorus removal are covered
in other sections of this manual and in several recent reports (1, 2, 3, 4).
6.1 Pre-Design Decisions
State quality standards on phosphorus usually form the major basis for considering
mineral addition for a given trickling filter plant. However, phosphorus concentration
in the effluent is only one factor involved; effluent BOD and suspended solids will also
be markedly reduced by this form of chemical treatment.
Modification of existing trickling filter plants for phosphorus removal is often a simple
operation, and inclusion of chemical treatment facilities in new plants is a relatively
minor addition. In either case, capital costs for chemical treatment equipment are
relatively low. Also, the required degree of phosphorus removal has little effect on
capital cost. Facilities for 80% phosphorus removal are identical to those for 95%
removal.
On the other hand, operating costs are substantially higher to obtain successful
treatment with chemical addition. More constant surveillance and more laboratory
support are needed than with a conventional plant. Added operating costs are on the
order of 4^ to 5^ per 1,000 gal. for reducing phosphorus (as P) to 2 mg/1 with an accom-
panying reduction of BOD and solids to 15 mg/1. For an additional 1^ to 2^ per 1,000
gal., phosphorus can be reduced to 1 mg/1 or less and BOD and solids to 10 or 12 mg/1.
This simultaneous reduction of phosphorus, suspended solids and BOD may be the
deciding factor for adopting chemical treatment in a trickling filter plant.
6.2 Process Options
There are three locations within a trickling filter plant where chemical precipitation
can be incorporated: the primary clarifier, the trickling filter, and the final clarifier.
The preferred approach is to provide facilities to serve both the first and last of these
but not to add chemicals directly to the trickling filter.
6 - 1
-------
Chemical addition to primary clarifiers will produce the greatest sludge yield of any
one arrangement. Since treatment at this point will reduce influent BOD about 50%, it
may be an appealing approach in an existing plant which is organically overloaded.
Unfortunately, a significant amount of the phosphorus present may not be in the
ortho form at this stage and may not be effectively precipitated and removed. Even
after chemical treatment, primary effluent may remain fairly turbid, but will clear up
during passage through the rest of the plant.
Mineral addition in the trickling filter does not seem to cause any functional problems
although the filter may blotch and slough more than usual. A modest inorganic film
often forms over the surface of the rocks but does not seem to interfere with
biological activity or cause ponding. As data shown later will indicate, however, high
degrees of phosphorus removal are not attained when chemicals are added only to the
filter. Poor removal may result partly from short circuiting in the filter. This approach
is not recommended.
Treatment in the final clarifier alone is a risky approach. If the operation is not done
effectively, poorly treated effluent can escape to the receiving water. In spite of these
reservations, experience has shown that precipitation in the finai clarifier can be both
effective and controllable. Orthophosphate predominates at this point and it
precipitates readily. Also, bio-degradable detergents which can interfere with
precipitation are largely absent. If underflow solids from a dosed final clarifier are
returned to the primary clarifier, it will stimulate unusually effective clarification there.
In order to allow for incorporating the advantages of chemical treatment in both the
primary and secondary clarifiers, it is recommended that chemical addition and mixing
facilities be provided at both locations. Facilities required are not expensive and
provision for mineral addition to both primary and secondary clarifier gives
considerable flexibility of operation. Chemical treatment in both clarifiers can then be
used if experience shows it to be the most effective technique at the particular plant
involved.
6.3 Performance Data Using Aluminum Salts
6.3.1 Richardson, Texas
A plant-scale study was begun at Richardson, Texas, in 1970 to evaluate the potential of
chemical addition to remove phosphorus from and improve overall performance of the
City's trickling filter plant (5). Major objectives were to reduce the effluent phosphorus
concentration to 1.0 mg/1 (as P), or less, and to reduce effluent BOD and SS residuals to 15
mg/1 or less.
The existing plant was a low-rate trickling filter facility with combination primary
clarifier/digester units (clarigesters) as shown on Figure 6-1. Operating data for a control
period before the addition of chemicals are shown in Table 6-1.
6-2
-------
FINAL SLUDGE
&
RECIRCULATION-
FINAL
EFFLUENT
INAL CLARIFIER
POLYMER
JUNCTION BOX
[TRICKLING
FILTERS
PRIMARY
CLARIFIERS
&
DIGESTERS
(UNHEATED)
SPLITTER BOX
SUPERNATANT
RAW
WASTEWATER
FIGURE 6-1 RICHARDSON, TEXAS
TREATMENT PLANT FACILITIES
DIGESTED
SLUDGE
.DRYING
BEDS
UPERNATANT
TREATMENT
6-3
-------
Table 6-1
OPERATING DATA FOR RICHARDSON, TEXAS
1970
Description Control Period
Flow, mgd
Influent BOD, mg/1
Influent SS, mg/1
Influent P, mg/1
Primary Clarifier
Overflow Rate, gpd/sq ft
Primary Effluent BOD, mg/1
Primary Effluent SS, mg/1
Primary Effluent P, mg/1
BOD Removal, percent
SS Removal, percent
P Removal, percent
Trickling Filter
Depth, ft
Recirculation Ratio
Hydraulic Loading, mgd/acre
Organic Loading, Ib BOD/ 1 ,000 cu ft/day
Secondary Clarifier
Alum Dosage Rate, Average Al :P mole ratio
Overflow Rate, gpd/sq ft
Secondary BOD Removal, percent
Secondary SS Removal, percent
Secondary P Removal, percent
Overall Plant Performance
Effluent BOD, mg/1
Effluent SS, mg/1
Effluent P, mg/1
BOD Removal, percent
SS Removal, percent
P Removal, percent
1.5
166
155
11
400
*
*
*
*
*
*
6.5
0
3.9
*
-
390
*
*
*
20
15
8
88
90
27
1971-72
Extended Alum Run
(11 % Months)
1.6
170
155
11.4
425
115
110
8.6
32
29
25
6.5
0
4.1
14.0
1.6**
415
96
94
94
5
7
0.5
97
95
96
*Not reported.
**Based on Influent P.
6-4
-------
Based on operational trials, alum addition ahead of the final clarifier was selected as the
most promising chemical additive system for an extended 1 P/z-month run. The flow diagram
for the upgraded system is shown on Figure 6-1.
It was found that a mole ratio (Al: Influent P) of 1.5 to 1.7 consistently yielded effluent
concentrations of 5 mg/1 of BOD, 7 mg/1 of SS and 0.5 mg/1 of total phosphorus (as P). The
corresponding effluent values for this plant prior to chemical addition were 20, 15 and 8
mg/1, respectively. The improved plant performance obtained with this upgrading technique
was attributed in part to the low final clarifier overflow rate, careful management of final
clarifier sludge withdrawal to prevent disruption to and loss of the alum floe blanket and
frequent manual adjustment of the chemical feed pump rate to match alum dosage to mass
inflow of phosphorus. Alum treatment doubled the volume of anaerobically digested sludge
produced. However, the digested alum/biological sludge exhibited superior drying charac-
teristics on sand beds and could be removed in about one-half the normal time. Operating
data for the upgraded plant are shown in Table 6-1.
Chemical costs were $0.05/1,000 gal of plant flow or $0.36/lb of phosphorus removed, with
phosphorus removal at the 96 percent level. The 1970 capital costs associated with plant
modifications for chemical addition were $65,000, allocated as follows:
New laboratory building $21,000
Laboratory equipment and furniture 6,000
Chemical storage and feed equipment, plant piping and
metering modifications 38,000
Total $65,000
6.3.2 Chapel Hill, North Carolina
Promising results were obtained in removing phosphorus and in generally improving plant
performance at Richardson, Texas, by the addition of alum to trickling filter effluent as
discussed in Subsection 4.5.4.1. In view of these results, the University of North Carolina
Wastewater Research Center initiated a follow-up study to further explore and confirm this
process at the Chapel Hill, North Carolina Wastewater Treatment Plant (6). Conducting a
similar chemical addition project at Chapel Hill was considered worthwhile because it is a
typical high-rate trickling filter plant using recirculation, whereas the Richardson filters were
low-rate units. Furthermore, parallel and identical lines of treatment units were available at
Chapel Hill, allowing direct comparison of results with and without alum addition.
Comparison of parallel results was not possible at Richardson.
The wastewater treatment plant at Chapel Hill is a conventional high-rate installation
treating predominately domestic wastewater. Incoming wastewater passes through pretreat-
ment facilities for the removal of large solids and grit. The flow is then divided into equal
portions for diversion to two identical lines of treatment, each consisting of a primary
6-5
-------
clarifier, trickling filter and final clarifier as shown on Figure 6-2. On one side of the plant,
a feed pump system was installed to add liquid alum to the trickling filter effluent just prior
to the final clarifier.
Operating data for the two sides of the plant are shown in Table 6-2. Significant
improvements in BOD, SS and phosphorus removals were achieved with an alum dosage rate
that ranged between an Al -.Influent P mole ratio of 1.5 and 2.2. Within this range and for
the loading rates shown, the final clarifier hydraulic loading appeared to be the significant
factor affecting process efficiency. Subsequent phases of the study indicated that
phosphorus removals of over 90 percent and BOD removals of about 95 percent could be
consistently obtained by lowering the final clarifier overflow rate to 500 gpd/sq ft.
Recirculation of the final clarifier settled alum/humus sludge to the primary clarifier for
thickening had a beneficial effect on primary treatment efficiency and decreased the organic
loading on the alum train trickling filter.
FIGURE 6-2
CHAPEL HILL, N.C. TREATMENT PLANT FACILITIES
PRIMARY TRICKLING FINAL
CLARIFIER FILTER CLARIFIER
INFLUENT
I
rO -
^^S
1 COMBINED
! SLUDGES
RECIRCU
LATION
1
> TFFLUENT
SETTLED FINAL
SLUDGE
SIDE 1(NO ALUM)
PRIMARY TRICKLING
CLARIFIER FILTER
INFLUENT
FINAL
CLARIFIER
COMBINED
SLUDGES
•»- EFFLUENT
RECIRCULATION
SETTLED FINAL
SLUDGE
SIDE 2(WITH ALUM)
6-6
-------
Table 6-2
OPERATING DATA FOR CHAPEL HILL, NORTH CAROLINA
Description
Flow, mgd
Influent BOD, mg/1
Influent SS, mg/1
Influent P, mg/1
Primary Clarifier
Overflow Rate, gpd/sq ft
Primary Effluent BOD, mg/1
Primary Effluent SS, mg/1
Primary Effluent P, mg/1
BOD Removal, percent
SS Removal, percent
P Removal, percent
Trickling Filter
Depth, ft
Recirculation Ratio
Hydraulic Loading, mgd/acre^
Organic Loading, Ib BOD/ 1,000 cu ft/day
Secondary Clarifier
Alum Dosage Rate, Average Al :P mole ratio
Overflow Rate, gpd/sq ft
Secondary BOD Removal, percent
Secondary SS Removal, percent
Secondary P Removal, percent
Overall Plant Performance
Effluent BOD, mg/1
Effluent SS, mg/1
Effluent P, mg/1
BOD Removal, percent
SS Removal, percent
P Removal, percent
Side 1
(No Alum)
1.41
168
229
11.3
1,100
77
89
9.7
54
61
14
4.25
2.0
16.3
18.8
—
885
43
29
5
44
63
9.2
74
72
19
Side 2
(With Alum)
1.41
168
229
11.3
1,100
63
68
6.6
63
70
42
4.25
2.0
16.3
15.4
1.722
885
76
53
73
15
32
1.8
91
86
84
^Includes recirculation.
^Based on Influent P.
6-7
-------
6.4 Performance Data Using Iron Salts
Addition of iron to trickling filters has been studied at Detroit and Wyoming, Michigan.
The Detroit Metro Water Department (7) compared the performance of pilot activated
sludge and plastic-media trickling filter systems in order to decide on the design basis
for a full-scale plant. Chemicals were added before the primary settler for both the
trickling filter and activated sludge studies, and to the primary effluent for the
activated sludge study. The plastic-media trickling filter test program consisted of a
series of 16 experimental runs. Application rates ranged from 1.0 to about 3.5
gpm/ft^ The raw wastewater inorganic phosphate averaged 7.9 mg/1 P with a range of
1.9 mg/1 to 26.6 mg/1. Without the addition of chemicals, total phosphate removals
ranged from 18 to 57%. With the addition of 15 mg/1 of iron (pickle liquor) removals
were 34 to 77%. With 15 mg/1 iron and 0.3 mg/1 Dow Purifloc A23 polymer, removals
ranged from 47 to 80%. With 15 mg/1 iron, 0.3 mg/1 A23 polymer, and 30 mg/1
NaOH, removals were in the range of 67 to 78%. It was concluded that control of
conditions for removal of phosphorus using the plastic-media trickling filters would be
more difficult than in the activated sludge process. The activated sludge process
achieved more consistent and higher phosphorus removals with iron addition and was
recommended for the full-scale plant.
Dow Chemical Company performed plant scale iron addition tests at the Wyoming,
Michigan trickling filter plant (8, 9, 10). The 1969 study was designed for ferric
chloride addition to the raw wastewater to take advantage of the additional BOD and
suspended solids removal in the primary. However, due to the recirculation of the
trickling filter sludge with its high phosphorus content, it was necessary to split the
FeClo addition between the raw wastewater and the trickling filter effluent. The
chemical feed and phosphorus removal data are presented in Table 6-3.
To achieve the 80% phosphorus removal required by the State, iron addition of 15
mg/1 to the raw wastewater and 3 to 4 mg/1 Fe to the trickling filter effluent
along with 0.3 mg/1 Purifloc A23 were necessary. Removal of BOD during chemical
treatment ranged from 60% to 87% across the plant. Suspended solids removal in the
primary averaged 75% with an overall removal of 87%. Iron addition to the Wyoming
trickling filter plant permitted greater than 80% phosphorus removal only with the
split chemical addition.
6.5 Choice of Chemical Addition
Since in-plant chemical precipitation is a new technique, pilot-scale data should be
obtained prior to design of new facilities. However, when the plant already exists,
full-scale trials may not be any more expensive than pilot-scale tests and they will be
more reliable.
6-8
-------
Table 6-3
IRON ADDITION FOR PHOSPHORUS REMOVAL AT WYOMING, MICHIGAN
Test No.
Chemical Feed
Fe to raw wastewater, mg/1
O i
Fe to trickling filter effluent, mg/1
Purifloc A23, mg/1, before primary
sedimentation
15
5-7
0.4
15
3-4
0.3
15-20
0
0.2-0.3
15
7
0
Total Phosphorus
Raw, mg/1
Primary effluent, mg/1
Removal, %
Final effluent,
Removal, %
mg/1
10.4
1.68
83.9
0.88
91.7
11.1
2.10
81.0
1.46
87.6
14.0
3.4
75.7
2.9
79.2
17.5
6.5
62.9
5.7
67.4
The jar test is a most important tool for both designer and operator when using
chemical addition. Information can be obtained inexpensively and in a short period of
time, however, certain precautions should be taken. An auxiliary flash mix is highly
desirable if the jar test apparatus does not have capability for true high energy mixing.
After intense dispersion of trial chemicals, flocculation can be simulated by careful
programming of time and mixing energy levels so they approximate conditions in the
plant. The most effective way to match jar test and plant conditions is to adjust the
jar test mechanism until it appears to duplicate the hydraulic regime observed directly
in the plant. Duration of each mixing interval should approximate plant conditions,
assuming plug flow. Finally, the machine should be left turning very slowly to approximate
settling conditions in the plant. A slight motion in the test jars is a better representation
of clarifier conditions than if stirring paddles are turned off. No matter how carefully
the rotation speed and timing intervals are selected to approximate plant conditions, the
operator should devote considerable practice to use of the apparatus before his results
are accepted and used.
A wide variety of other bench scale and laboratory tests (11, 12) hold considerable
promise in wastewater technology. At this time, however, none of them have proved as
effective as the jar test.
6.6 Nature and Role of Chemicals Involved
Compounds mentioned here are those most likely to be selected for chemical
treatment in a trickling filter plant. Of the commercially available metal salts (13),
liquid forms are more convenient to handle, especially at small plants. They may have
6-9
-------
the lowest overall cost and are generally more effective than dry forms. However, it
may be necessary to use dry salts because transportation of pre-mixed solutions,
containing both water and salt, may be costly.
Metal salt forms include liquid ferric chloride, pickle liquor, liquid alum, and sodium
aluminate. Information on properties and handling these chemicals is given in Chapter 10.
Polymers are available in a variety of forms, mostly dry powders. Since the amount of
polymer required is far less than the amount of metal salt, the problems of using dry
polymer compounds are not severe. Of the three categories of polymers (cationic,
anionic and nonionic) there is no universal choice regarding the most effective type.
Whenever chemicals are chosen, they must perform two main functions: precipitate
phosphorus, and remove all types of colloids in the water. Phosphorus precipitation is
caused solely by metal salts and involves conversion of soluble orthophosphate to
insoluble phosphate colloids. Following that, both metal salts and polymers serve to
coagulate (destabilize) colloids. Flocculation of destabilized colloids is also brought
about by both metal and polymers.
6.7 Dosage Selection and Control
Proper dosage selection and control are major keys to the success of chemical
treatment because this is where both cost and performance are determined. The
technique selected must be effective but also should be simple and easily interpreted
by the operator.
In the metal salt system, the key parameters are the mole ratio fed and the effluent
phosphorus concentration. As a basis for dosage determination, complete composite
hourly and daily phosphorus analyses should be performed initially. These data will
provide a guide to both total daily dose required and to daily variations in dosing rate.
Once accomplished, this task need not be repeated because subsequent refinements can
be based on phosphorus concentration in the final effluent. Phosphorus observations
should be made at both the primary clarifier and the final sedimentation tank if direct
dosing of the filter is discounted.
A suggested procedure for setting the dose is as follows. First, express incoming
phosphorus in pounds per day, then convert to Ib-mole/day by dividing by the atomic
weight (Ib/lb-mole). Atomic weights are 31 for phosphorus, 27 for aluminum and 56
for iron. Next set the mole ratio (metalrP) at the proper value, between 1.5:1 and
2:1, and compute the daily coagulant dose. An example calculation is shown below:
If incoming phosphorus (as P) is 310 Ib/day,
310/31 = 10 Ib-moles/day
If desired mole ratio (M:P) is 2:1,
2
(-) (10) = 20 Ib-moles metal required
6-10
-------
Using liquid alum having 4.37% Al,
(20) (27) = 540 Ib Al required
(540)/(11.1 Ib/gal.) (0.0437)
= 1,100 gal. liquid alum required/day
Then take the total volume to be fed per day and convert it into dosage rates with
three to five changes of rate per day. These different feed rates should be selected to
match incoming phosphorus at the point of injection. Dose must be matched carefully
to phosphorus demand. Feeding twice the amount required part of the time and half
the amount the remainder of the time will waste chemicals and lead to poor
performance. In fact, overfeeding chemicals can give as poor results as underfeeding;
this stems from the tendency of high metal concentrations to drive phosphorus colloids
into a highly stabilized state. Dilution of sewage during rainy weather makes little
difference in the amount of metal salts required since total phosphorus, in terms of Ib/day,
remains about the same.
Changes in feed pump settings can be made manually according to a schedule
established and posted beforetime. Several types of cam regulated feed controls are
also available and are effective (14).
After establishing the initial feed program, effluent phosphorus should be monitored
hourly or continuously. The plant should be kept on a given feed schedule for at least
three to five days, to allow the biological reactions to adjust to any effects of
chemical dosing. The effluent phosphorus data will show peak concentrations resulting
from dosing improperly for a period of time. The feed schedule can then be adjusted
to remove these peaks. Finally, studies of different mole ratios fed can be made and
the results plotted as shown on Figure 6-3. These curves can then be used to choose
the mole ratio required to meet effluent phosphorus standards.
Continuing operation of the plant "should be monitored carefully and adjustments
should be made in the chemical feed schedule when indicated by consistently high
effluent phosphorus. Chemical feeding can be automated.
Polymer may be needed for good solids removal. Control of polymer feeding is
relatively straightforward. Polymer feed rates are usually kept constant, but a reduced
feed rate may be used during periods of low flow. Use of polymers may reduce metal
salt demand and make metal addition more attractive. Polymers which prove effective
in jar tests should be carefully evaluated in plant scale operation. Typical dosage is 1
mg/1 or less. One direct way to judge polymer effectiveness is to observe effluent
turbidity. Polymer feeding can also be automated.
6-11
-------
d SB I/BUI
HI snuoHdsoHd ivioi
co
<
X
LLJ
00
oooo_
00
CO
-------
6.8 Iron Leakage
Iron leakage may be a problem in trickling filter plants using iron precipitation. Figure
6-4 shows some typical data from Richardson, Texas (5). The problem appeared worse
when iron was being fed in the final clarifier. It is evident that polymer treatment will
reduce the escape of iron colloids.
6.9 Sampling and Analysis
The use of chemicals to remove phosphorus in trickling filter plants will require greater
laboratory space, facilities, and reagents than normal treatment. Table 6-4 shows a
typical matrix of data which is desirable at a conventional plant, as well as typical
analyses required to support chemical treatment.
In the latter matrix, notice that the analyses include both anion and cation of the
chemical being fed. Monitoring of alkalinity is optional unless the effluent
concentration drops below 50 mg/1 as CaCOo. Some alkalinity is required to provide
buffering and prevent a lowering of pH which would have an adverse effect on both
chemical precipitation and biological activity.
Observation of turbidity in the final effluent is a good test. There is strong evidence
that phosphorus and turbidity correlate in final effluents and the turbidity test
provides a simple and responsive on-site control tool. Turbidity may be measured in
the laboratory with a bench top meter or in the plant with a submersible disk lowered
in the final clarifier. Observation of the height of the floe blanket in the clarifier is
also a good means of control.
Effluent phosphorus concentration should be determined either on a semicontinuous
basis with an automatic analyzer or on an hourly basis using automatic or manual
sampling. In addition, a daily composite is also advisable. Sample portions should be
flow-weighted rather than time-weighted. In plants enjoying stable effective
performance, the schedule may be relaxed.
Signals from direct reading sensors and automatic analyzers can be used as a basis for
automatic control of plant operation.
6-13
-------
10
LU
(a) IRON LEAKAGE IN PLANT EFFLUENT
WHEN TREATING IN
FINAL CLARIFIER
. .WHEN TREATING
/PLANT INFLOW
I
I I I I
I I
0.8 1.0 1.2 1.4 1.6 1.8 2.0 2.2 2.4
MOLE RATIO, Fe3+/P
0
0.8
(b) IRON LEAKAGE WHEN TREATING PLANT INFLOW
IRON LOST IN
EFFLUENT
W/0 POLYMER
I
RON LOST IN EFFLUENT
W/POLYMER IN FINAL
CLARIFIER
I
1.0 1.2 1.4 1.6 1.8 2.0
MOLE RATIO, Fe3+/P
2.2 2.4
FIGURE 6-4 IRON LEAKAGE DURING PHOSPHORUS REMOVAL
6-14
-------
TABLE 6-4 DESIRABLE LABORATORY ANALYSES
ANALYSES FOR CONVENTIONAL TREATMENT
FLOW
TOT SOL
TOT VOL SOL
SUS SOL
SUS VOL SOL
SET SOL
BOD
DO
COD
pH
TEMP
R
A
W
X
X
X
X
X
X
X
X |
X
X
X
p
R
1
M
E
F
F
X
X
X
X
X
X
X
X
X
F
1
L
T
E
F
F
X
X
X
X
X
X
X
X
F
1
N
A
L
F
F
F
X
X
X
X
X
X
X
X
X
X
R
E
C
1
R
C
X
X
X
X
X
X
S
L
U
D
G
E
X
X
X
X
X
S
U
p
E
R
N
A
T
X
X
X
X
X
X
X
X
X
ADDITIONAL ANALYSES FOR CHEMICAL TREATMENT
PHOS
ALK
FE
AL
so4
CL
TURB
R
A
W
X
X
X
X
X
X
X
P
R
1
M
E
F
F
X
X
X
X
X
X
X
F
1
L
T
E
F
F
X
F
1
N
A
L
E
F
F
X
X
X
X
X
X
X
R
E
C
1
R
C
X
X
X
S
L
U
D
G
E
X
X
X
X
X
S
U
P
E
R
N
A
T
X
X
X
X
X
6-15
-------
6.10 References
1. Nesbitt, J. B., "Phosphorus Removal - The State-of-the-Art", JWPCF, 41:5, p
701 (1969).
2. Hall, M. W., and Engelbrecht, R. S., "Phosphorus Removal - Past, Present, and
Future", Wat. and Wastes Eng., 6:8, p 50 (1969).
3. Scalf, M. R., et al, "Phosphate Removal: Summary of Papers", Jour. SED, ASCE,
95:SA5, p 817 (1969).
4. Stephan, D. G., and Schaffer, R. B., "Wastewater Treatment and Renovation -
Status of Process Development", JWPCF, 42:3, p 399 (1970).
5. Derrington, R. E., et al., "Enhancing Trickling Filter Performance by Chemical
Precipitation," U.S. EPA Project No. 11010 EGL.
6. Preliminary Report, Methods for Improvement of Trickling Filter Plant Performance.
Part I, University of North Carolina, U.S. EPA Contract 14-12-505.
7. Detroit Metro Water Department, "Development of Phosphate Removal Processes",
Detroit, Michigan (July, 1970) (Program 17010 FAH Grant WPRD 51-01-67
Advanced Waste Treatment Laboratory, Cincinnati, Ohio), Pre-publication copy.
8. Condon, W. R., "Design of the Wyoming, Michigan Wastewater Treatment Plant
Improvements", presented at the Technical Seminar/Workshop on Advanced Waste
Treatment, Chapel Hill, North Carolina (Feb. 9-10, 1971).
9. "Report on Wastewater Treatment Plant Improvements, Wyoming, Michigan", Black
and Veatch of Michigan, Consulting Engineers, Kansas City, Missouri (1970).
10. Stonebrook, W. J., Dykhuizen, V., Beeghly, J. H., and Pawlak, T. J., "Phosphorus
Removal at a Trickling Filter Plant, Wyoming, Michigan", Unpublished (undated).
11. TeKippe, R. J. and Ham, R. K., "Coagulation Testing: A Comparison of
Techniques - Part 1", JAWWA, 62:9, p 594 (1970).
12. TeKippe, R. J. and Ham, R. K., "Coagulation Testing: A Comparison of
Techniques - Part 2", JAWWA, 62:10, p 620 (1970).
13. "Coagulants for Waste Water Treatment", Chem. Eng. Prog., 66:1, p 36 (1970).
14. McAchran, G. E. and Hogue, R. D., "Phosphate Removal from Municipal
Sewage", Wat. & Sew. Works, 118:2, p 36 (1971).
6-16
-------
Chapter 7
PHOSPHORUS REMOVAL IN ACTIVATED SLUDGE PLANTS
BY MINERAL ADDITION
7.1 Description of Process
The mechanism of phosphorus removal by mineral salt addition, usually iron or
aluminum compounds, in a biological system is through a combination of precipitation,
adsorption, exchange, and agglomeration as influenced by the pH and ionic composition
of the water. The phosphorus removal technique is operationally simple and is
accomplished by direct mineral addition to an aeration tank. Treatment costs are largely
a function of the required effluent phosphate residual. Through optimization techniques,
any degree of phosphorus removal may be provided. The main liability is the
introduction of dissolved solids.
Phosphorus removal by biological systems alone, regardless of the internal mechanisms of
phosphorus incorporation into the sludge solids, is limited, and removal is primarily
controlled by the magnitude of solids wasting. The net amount of solids wasting is
dictated by the quantity and type of substrate received and removed, system design and
operation, and solids capture in the final effluent. A high rate system, with its greater
solids production, would undoubtedly have the greatest background phosphorus removal
potential by pure synthesis mechanisms. The magnitude of this background removal can
be found by plant measurement with existing facilities or estimated by knowledge of the
net solids production and phosphorus concentration in the waste activated sludge. Figure
7-1 shows the phosphorus removal capability for a variety of activated sludge systems at
typically reported values for domestic wastewater treatment.
Historically, the benefits derived from mineral additions to activated sludge plants have
been known since 1938 (1). Thomas (2) was among the first to point out the
increased phosphorus removals. Wirts (3) also observed the increased phosphorus
removals derived from a metal bearing industrial waste at a municipal treatment facility
in Cleveland, Ohio. Early investigators of mineral addition for upgrading and modifying
activated sludge systems for phosphorus removal include Tenny and Stumm (4), Earth
and Ettinger (5), and Eberhardt and Nesbitt (6). Earth, et al., conducted pilot plant
investigations with a three sludge system for combined chemical-biological nitrogen and
phosphorus removal (7). Large scale, long term investigations have been recently
completed at Manassas, Virginia (8), Pennsylvania State University (9), and Texas City,
Texas (10). Recent publications have examined the use of aluminum (11) and iron
(12) salts and investigated the kinetics and mechanisms of phosphorus removal using
these minerals (13).
In the treatment of domestic wastewaters several investigators, most recently Rickert
and Hunter (14), have shown that in a single sludge system a stable organic residual is
found after aeration times of 30 to 60 minutes. Unfortunately, if treatment is stopped
7 - 1
-------
0.5
O.U -
0.3
0.2 -
I
TYPICAL
PERFORMANCE RANGE
0.03 P/VSS
HIGH RATE
A. 8.
STD. RATE
A.S.
EXT. AER.
A.S.
Assume COD Removal = 200 mg/1
(Prrmary Effluent - Soluble Final Effluent)
I
01234
PHOSPHORUS REMOVED, rag/I as P
FIGURE 7-1 PHOSPHORUS REMOVAL CAPABILITY OF ACTIVATED
SLUDGE SYSTEMS RECEIVING SETTLED DOMESTIC WASTEWATER
7 - 2
-------
at this point, bio-flocculation is almost nonexistent and poor solids densities, settling
characteristics and effluent quality result. However, when incorporating mineral
addition, these liabilities are overcome and there exists an opportunity to derive capital
savings from reduced tankage requirements to compensate for the increased operating
expenditures associated with mineral addition. Indeed, in many cases, the use of a
short residence time activated sludge system with mineral addition, settling, and
filtration may result in a higher quality effluent for equal or lower costs than a purely
physical-chemical approach of coagulation, settling, filtration and carbon adsorption
(15). From a liability standpoint, however, the designer should be cognizant of the
greater mass of generated waste activated sludge which must be successfully handled
and treated for ultimate disposal.
Mineral addition to activated sludge for phosphorus removal offers the following
advantages:
a. Ease of operation (direct chemical addition to existing unit processes).
b. Relatively small additional solids production (causing increases in sludge
density and dewaterability).
c. Flexibility to changing conditions (treatment costs are largely a function
of the influent and required final effluent phosphorus concentrations).
The main disadvantage is the introduction of dissolved solids which in some
circumstances can be considered pollutants in their own right.
7.2 Mineral Selection and Addition
The selection of the mineral to use for phosphorus removal involves a multitude of
considerations. Any commercial chemical, or industrial waste stream (for example,
pickle liquor or alum sludge from a water treatment plant) bearing available aluminum
or iron has potential application. Ideally, the best approach requires comparative
full-scale plant tests of several months duration. This, of course, is usually impossible,
and the next best approach is to run comparative jar tests, the results of which may
not be directly applicable. However, they are at least indicative of performance. As an
illustrative example, the procedures and results from the Manassas studies (8, 16) are
utilized in the subsequent paragraphs.
Figure 7-2 shows the total soluble phosphorus as a function of the metal ion dose for
the Manassas raw wastewater, for mixed liquor from a high rate activated sludge
system, and for biologically stabilized final effluent. Since both the alum and ferric
chloride can cause an alkalinity depletion and potential pH drop, pH values are also
recorded. Superior performance for alum was found at all mass dosages with an
apparent limiting residual of 1.0 mg/1 as P for the raw wastewater and final effluent,
and less than 0.1 mg/1 as P for the aerator mixed liquor. A yellow color in the
product water was noted with the higher iron dosages.
Replotting the aluminum data from Figure 7-2 in Figure 7-3 shows the influence of
7 - 3
-------
FERRIC CHLORIDE
AS Fe3±
20
PH
Fe3+
A13+
6.9
6.9
6.9
6.7
6.8
6,5
6.7
6.1
6.6
5.9
6.5
01
e
Q_
c/o
o
O
Cf>
Fe
3+
AERATOR
HIGH RATE ACTIVATED SLUDGE
MLSS = 1850 mg/l,%VOL= 80
10
20 30 40
METAL ION DOSE, mg/1
50
FIGURE 7-2 TOTAL SOLUBLE PHOSPHORUS RESIDUAL VERSUS
METAL ION DOSE
(Reprinted with permission from the American Institute
of Chemical Engineers, New York, N.Y. 10017)
7 - 4
-------
1.2
1.0
0.8
CO
O
O
oo
«t
I—
O
O
0.6
O.U
0.2
RAW WASTEWATER
I
1.0 2.0 3-0 4.0
MOLES ALUMINUM ADDED
MOLE INITIAL TOTAL SOLUBLE PHOSPHORUS
5.0
FIGURE 7-3 INFLUENCE OF POINT OF ADDITION
ON PHOSPHORUS REMOVAL WITH ALUMINUM
(Reprinted with permission from the American Institute
of Chemical Engineers, New York, N.Y. 10017)
7 - 5
-------
the point of addition on the molar effectiveness of the added aluminum. Similar but
lower molar curves are found with the iron versus phosphorus residual data. The lag in
phosphorus removal effectiveness encountered with the raw wastewater and aerator
mixed liquor is thought to be due to the presence of unhydrolyzed phosphorus and
the occurrence of competing side reactions for the metal ion. Aluminum and iron are
known to form insoluble heavy salts with certain hydrophilic colloids such as
detergents and soluble proteins and their degradation products (17), all of which would
be present in raw and partially stabilized wastewater. These curves indicate that at
molar ratios of less than 1:1, there is an obvious advantage in phosphorus removal
derived by adding the chemical to a well stabilized liquid, whereas at molar ratios
above 1.5:1, enough metal ion has been added to satisfy the competing side reactions
and all points of addition are equally effective. The data indicate that very low
phosphorus residuals are possible in a combined chemical-biological system and that the
ideal point of addition would be just before final solids-liquid separation to obtain
maximum biological stabilization of the soluble liquid phase.
Figure 7-4 shows the influence of wastewater and process pH upon phosphorus
removal effectiveness with and without mineral additions. In the upper graph of Figure
7-4, above pH 8, phosphorus removal is occurring naturally through the formation of
an insoluble calcium-phosphate complex via the natural hardness of the wastewater.
The lower graph of Figure 7-4 shows that phosphorus removal at any molar dosage of
metal ion is actually controlled by pH dependent reactions, optimally around 6 and 5
for aluminum and iron, respectively. Since the optimum range for aerobic oxidation of
carbonaceous compounds is about pH 6 to 8, it can be seen that there is an
opportunity for the simultaneous optimization of carbon oxidation and phosphorus
removal by aluminum addition. However, if nitrification is a treatment goal there exists
a conflict and, in turn, a compromise in pH goals. Diminished nitrification performance
has been reported when alum additions have depleted the wastewater's alkalinity to a
point where the pH was near optimum for phosphorus removal (9).
Table 7-1 provides a further comparison of representative sources of aluminum and
iron. It should be noted that dissolved solids or extraneous ions are introduced and in
some circumstances these ions may be undesirable (18). Alum and FeCl3 cause an
alkalinity loss, and without buffering capacity in the wastewater, a substantial pH
reduction may occur. This can be compensated for by adding a source of alkalinity,
such as lime, or by using aluminate. Treatment performance can deteriorate
dramatically if the system pH is allowed to fall much below optimum (5). The
designer should be aware that periods of heavy storm water infiltration will
substantially dilute the normal alkalinity of the wastewater, and if operation is carried
out at a pH of 6, only 20 to 40 mg/1 of residual CaCO3 buffering alkalinity may
remain (8). The bottom part of Table 7-1 shows a comparison of the metal salts on a
mole per mole basis, assuming a mole of metal ion precipitates one mole of
phosphorus. In practice this never occurs. On this basis the possible cost advantage and
alkalinity advantage with iron is lost and the aluminum salts show a potential for less
additional solids production.
7 - 6
-------
o.
oo
o
o
oo
•a.
I—
o
20
10
*
, ,
AEROBIC BIOLOGICAL
TREATMENT ZONE
ZONE FOR
PHOSPHORUS
REMOVAL
h *
ZONE FOR
HITRIF1CATIOI
* *
\
I
\ T l\
NO METAL
ION ADDED
7
pH
10
o
5
oe
Q.
V)
O
o
00
00
LLJ
1.00
0.80
I3+ ADDED AS ALUM
Fe3+ ADDED AS FERRIC CHLORIDE
I Mole Metal Ion
CHEMICAL DOSE:
Mole Initial Total Soluble Phosphorus
i i i
0.20
10
FIGURE 7-U EFFECT OF pH ON PHOSPHORUS REMOVAL FROM A
FINAL EFFLUENT
(Reprinted with permission from the American Institute
of Chemical Engineers, New York, N.Y. 10017)
7 - 7
-------
Table 7-1
METAL SALT COMPARISON
Metal Salt
Liquid
Sodium
Aluminate
Liquid
Alum
Ferric
Chloride
(A13+) (A13+) (Fe3+)
(per Ib of Metal Ion)
Bulk Delivered Price at Manassas, $
Extraneous Ion
Introduced, Ib
Maximum Alkalinity Loss as CaCOo, Ib
0.69
Na+
0.85
None
0.32
S042'
5.4
5.5
0.23
cr
1.9
2.7
(Per Ib of Phosphorus at a Mole Dosage of 1::
Bulk Delivered Price at Manassas, $
Extraneous Ion
Introduced, Ib
Maximum Alkalinity Loss as CaCOo, Ib
Phosphorus Removal of Metal Ion, Ib
Additional Solids Production, Ib
0.60
Na+
0.74
None
0.87
0.28
S042"
4.7
4.8
0.87
0.41
cr
3.4
4.8
1.8
AlPO^
3.9
AlPO^
3.9
FePO^
4.9
Several other points should be considered in mineral selection. First, in the process
effluent the soluble Al, less than 0.5 mg/1 (16), is apparently much lower than found
with iron additions, 6 mg/1 Fe reported at Cleveland (19). Second, alum adds a
divalent ion (SO4 ) which has a proven coagulation benefit (20). Third, both
Mulbarger (8) and Long, et al. (9), report improved product water clarity when using
alum in comparison to aluminate.
Chemicals can be added with confidence in the activated sludge process just before
solids-liquid separation, assuming adequate dispersal. Studies by Allied Chemical
Company (21) in a 25,000 gpd pilot plant indicated that the addition of alum at the
seven-eights point in the aeration basin, along with anionic polymer in the transfer
line, provided the optimum phosphorus removal, coagulation and settling. Polymer
dosages were found to be optimum in the 0.2 to 0.4 mg/1 range. An average A1:P
weight ratio of 2.1:1 was required to reduce the average influent total phosphorus of
3.6 mg/1 to 0.45 mg/1 in the effluent.
It should be noted that utilization of activated sludge aeration tanks for chemical
7 - 8
-------
mixing and flocculation is a compromise. Typically, a diffused aeration system will
have a velocity gradient of over 100 ft/second/ft and the velocity gradient of a
mechanical aeration system will be much higher. Water treatment experience has shown
that velocity gradients greater than 75 ft/second/ft will result in the onset of chemical
floe disintegration (22). Undoubtedly, some chemical floe as well as biological floe
disintegration does occur in activated sludge tanks and, because of solids recycle, is
unavoidable. Floe disintegration and mineral overdose both lead to product water
quality degradation. In new design situations it may be advantageous to provide a brief
period of gentle solids agitation prior to solids separation to promote reflocculation of
the chemical-biological solids. This can be done by gentle air agitation, a flocculation
tank, or a flocculator-clarifier. The designer should also provide the capability of
adding organic polymer, if needed. Nalco 676 and Atlas 2A2, slightly anionic organic
polyelectrolytes, have been successfully used at dosages between 0.3 and 0.4 mg/1 (8).
7.3 Performance and Optimization
7.3.1 Phosphorus Removal
At molar dosages of 1 to 2, metal ion to phosphorus, soluble phosphorus residuals of
0.3 to 3.0 mg/1 as P are frequently encountered in the literature. Interpretation of
some of the data is complicated by failures to establish biological equilibrium [90%
equilibrium values are reached at approximately 2.2 times the operating cell retention
time (16)] prior to changing variables or initiating an investigative run, and failure to
consider and report pH.
Detailed data from iron additions to activated sludge systems are limited in the
literature and consequently the more extensively available data on aluminum addition
provide a better source of definitive design information. However, a few investigations
of iron additions to secondary treatment processes have been conducted on the pilot
and full plant scale.
Pilot plant studies by the Detroit Metro Water Department, in which several types of
activated sludge systems were investigated, indicated that the most effective total
phosphate reduction was provided by the conventional activated sludge process with
iron (FeC^ derived from pickle liquor) fed to the primary effluent (23). The majority
of the test runs exhibited removals varying from 75 to 85% for inorganic phosphate
concentrations ranging from 1.9 to 26.6 mg/1 as P in the raw wastewater. Phosphate
removals by the activated sludge process alone (no iron fed) fell principally in the 55
to 70% range. It was found that, in general, the addition of iron to the conventional
activated sludge process provided a treatment system capable of consistently high
phosphorus reduction. The step-feed, activated sludge process, with the addition of iron
or iron, polymer and caustic soda, also provided consistent total phosphorus reduction.
It was concluded that the full-scale plant should be designed for the activated sludge
process, arranged for both conventional and step-feed configurations, and that provision
be made for phosphorus removal, utilizing injection of steel pickle liquor
7 - 9
-------
A pilot plant (5,760 gpd) at Warren, Michigan (24) was designed to simulate the existing 36
mgd activated sludge plant, with the addition of one sand filter and one multimedia
filter operated in parallel after the final clarifiers. Ferric chloride was fed to the
effluent of the aeration basins to determine whether or not this method would meet
the State-required 80% phosphorus removal. Ferric chloride addition at a ratio of 1.5
mg/1 of Fe to 1.0 mg/1 of soluble PO^ reduced the average total phosphorus
concentration from 7.0 mg/1 to 1.2 mg/1 in the effluent. This amounted to a
reduction of 83%.
Leary, et al., of the Milwaukee Sewerage Commission (25), conducted a one-year
plant-scale study of the use of waste pickle liquor (FeSO^) to enhance phosphorus
removal. The Milwaukee Jones Island Plant consists of a single primary facility
followed by two parallel activated sludge, plants. Pickle liquor was employed for
phosphorus removal at the 115 mgd East plant during the test period. The 85 mgd
O-i_
West plant served as a control. Pickle liquor feed was equivalent to an average Fe
dosage of 9.4 mg/1. The pickle liquor was added before the aeration tank about
55 feet downstream from the point of addition of return sludge. Based on a 1970
average total phosphorus concentration of 8.2 mg/1 P in the screened sewage, the East
Side plant, with iron addition, accomplished 91.3% phosphorus removal while the
control West plant removed 83.1%. Effluent phosphorus averaged 0.7 mg/1 at the East
plant and 1.4 mg/1 at the control plant. The return sludge phosphorous concentration
was 2.29% P in the control plant and 2.61% P in the East (test) plant. The iron
content of the return sludge was increased from 1.86 to 5.08%.
2_
The pickle liquor addition increased the East plant effluent 804 concentration from
123 to 145 mg/1 and decreased the alkalinity from 213 to 169 mg/1 as CaCO3-
Yearly average pH values were 7.0 and 7.1, respectively, for the East and West plants.
The free acid in the pickle liquor used (two sources) ranged from 2.1 to 9.3%
H2S04.
Comparison of the BOD, COD, and suspended solids removal efficiencies of the West
and East plants indicated that the addition of unneutralized pickle liquor did not
adversely affect purification, and it did not cause problems with the plant physical
facilities. The results of this one-year study indicated that the use of waste sulfuric
acid pickle liquor as a source of iron for phosphorus precipitation and removal was
practical and effective in maintaining low plant effluent phosphorus residuals.
In studies at the University of Missouri at Rolla (26), alum and aluminate at various
times were fed to laboratory activated sludge units and compared to a control unit.
The alum was more effective, as shown on Figure 7-5, than aluminate, based on the
A1:P mole ratio. In the activated sludge units which were dosed with aluminum, the
volatile mixed liquor suspended solids dropped to 50 to 60% of the total mixed liquor
suspended solids, compared to a value of 80% in the control. The aluminate units had
higher pH, but all pH values were in the range of 7.7 to 8.2.
7 - 10
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Full-scale studies were conducted at Pennsylvania State University on the removal of
phosphorus by addition of aluminum to a 2.0 mgd activated sludge plant. These studies
began in 1969 and continued through 1970 (9). In the first phase of the work, alum
and sodium aluminate were compared with respect to their ability to precipitate
phosphorus. To prevent biological upset due to mineral solids displacing biological
solids in the mixed liquor, the volatile suspended solids concentration was held at the
same level. Since the volatile fraction of the total mixed liquor suspended solids
decreased due to mineral solids buildup, this required the maintenance of a higher
total suspended solids concentration in the mixed liquor.
The data shown by Table 7-2 indicate that alum was slightly superior to sodium
aluminate in removing phosphorus from the moderately alkaline wastewater. The best
point of addition for alum was found to be at the effluent channel of the aeration
basin which carries mixed liquor to the final settling basin. Addition of alum both at
the inlet end of the aerator and at a point about two-thirds of the way along the
basin from the inlet resulted in lower removal efficiencies. Alum addition ahead of
aeration produced a cloudy effluent.
Table 7-2
COMPARISON OF USE OF ALUM AND
SODIUM ALUMINATE AT PENNSYLVANIA STATE UNIVERSITY
Chemical*
Parameter Alum Sodium Aluminate
Influent Effluent Influent Effluent
Suspended Solids, mg/1 134 19 94 25
BOD5, mg/1 87 6 50 7
COD, mg/1 244 31 160 36
Phosphorus as P, mg/1
Filtered Ortho 6.4 0.28 5.4 0.51
Filtered Total 7.6 0.27 6.9 0.57
Unfiltered Ortho 10.4 1.17 6.2 1.14
Unfiltered Total 12.4 1.48 8.6 1.53
Alkalinity as CaCO3, mg/1 156 103
pH 7.4 7.1 7.15 7.25
Alum Dosage as A12(SO4)3- 14H2O, mg/1 160
Na2Al2O4 dosage, mg/1 44
A1:P Ratio **, weight 2.04:1 2.04:1
S042-, mg/1 24 131
*Chemicals added at the effluent end of the aeration tank.
**A1:P ratio based on expression of phosphorus in the filtered ortho form.
7 - 12
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Aluminate addition was found to be most effective at a point in the aeration basin
near the outlet. Aluminate addition either before or after aeration produced a more
turbid effluent.
Sludge production in the studies comparing alum and aluminate showed that the
volume of sludge produced per million gallons of wastewater treated was about the
same for both coagulants. However, the weight of solids produced was considerably
less for aluminate than alum.
Because alum was found to be superior to aluminate in removing phosphorus, further
study of alum was performed for a period of one year. One-half of the activated
sludge plant was used in the alum study and the other half was operated without
chemical addition for comparison. Table 7-3 shows average results for operation at
rates equal to or less than the plant design capacity.
Table 7-3
COMPARISON OF ALUM ADDITION TO THE AERATOR AND CONVENTIONAL
ACTIVATED SLUDGE AT PENNSYLVANIA STATE UNIVERSITY
FOR FLOWS NOT EXCEEDING THE DESIGN CAPACITY OF THE PLANT
Effluent with Effluent without
Parameter Influent Alum Addition* Alum Addition
Suspended Solids, mg/1 110 22 26
BOD5,mg/l 71 9 13
COD, mg/1 172 55 68
Phosphorus as P, mg/1
Filtered Ortho 6.7 0-28 6.7
Filtered Total 6.3 0.36 6.7
Unfiltered Ortho 8.9 1.15 7.2
Unfiltered Total 10.0 1.41 7.3
Alkalinity as CaCO3, mg/1 168 80 120
pH 7.6 6.75 7.35
A1:P Ratio**, Weight 2.97:1
*Alum added to effluent of aeration tank. Values shown in this column are means
for the period of study.
**The A1:P ratio is based on expression of phosphorus in the filtered ortho form.
Problems were experienced with loss of solids from the final basin during peak flow
periods where alum treatment was being practiced. It was believed, however, that
proper allowance for peak flows in design of such facilities would eliminate excessive
solids loss. The effluent insoluble phosphorus concentration was found to correlate well
with the effluent suspended solids concentration.
A 2,000 gpd pilot operation at the main wastewater treatment plant of the District of
7 - 13
-------
Columbia was operated at a constant alum dose rate (27). The pilot plant is a
four-stage step aeration process followed by secondary clarification. Alum was added to
the wastewater of moderate alkalinity (100 to 150 mg/1 as CaCO-j) at the fourth stage
of aeration using an A1:P weight ratio of 2:1 (molar ratio 2.3:1), based on an influent
phosphorus concentration of 8.15 mg/1 as P. Mixed liquor suspended solids averaged
6,300 mg/1 with a volatile content of 60% and a cell retention time of about 5 days.
Total phosphorus residuals averaged 1.9 mg/1 as P (78% removal) in the clarifier
effluent.
Figure 7-6 shows the chemical phosphorus removal found for the Manassas high rate
activated sludge plant with aluminate (16) and alum (8) additions at slightly alkaline
pH values. Phosphorus concentrations for Figure 7-6 were corrected for biological
removal. The Manassas Pilot Plant, shown in Figure 7-7, is a 0.2 mgd three-stage
(nitrification-denitrification) activated sludge system operated at the Greater Manassas
Sanitary District of Prince William County, Virginia. Table 7-4 shows the system
performance for 200 determinations with alum addition. The reader is referred to the
original publication for more detailed information.
Table 7-4
COMBINED CHEMICAL-HIGH RATE
ACTIVATED SLUDGE SYSTEM PERFORMANCE
Percent Less Than
Item Min. _10 20 50 80 90 Max.
Primary Effluent
COD, mg/1 27 64 86 152 227 250 324
P, mg/1 2.2 3.8 5.0 8.4 11.4 12.8 18.5
BOD5:COD* 0.26 0.43 0.56
High Rate Activated Sludge Effluent
Soluble Phase
COD, mg/1 13 23 26 33 41 48 79
P, mg/1 0.2 0.8 1.0 2.0 3.0 3.6 6.8
BOD5:COD* 0.14 0.28 0.36
Solid Phase
Suspended Solids, mg/1 0 10 14 21 36 44 196
BOD5:SS* 0.10 0.23 0.31
COD:SS* 0.43 0.58 0.80
P:SS* 0.039 0.051 0.063
Suspended Solids After
Polymer Addition,** mg/1 00 1 2 8 1418
Suspended Solids After
Mixed Media Filtration**, mg/1 000 0 2 5 12
*Ratio of minimum, median, and maximum monthly averages for each parameter
**94 determinations
7 - 14
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7 - 16
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Optimization of phosphorus removals with mineral additions can be done two ways:
by pH adjustment to the point of minimum solubility (pH 6 to 6.5, residual alkalinity
of 20 to 40 mg/1 as CaCOo) and by splitting the chemical dose to separate parts of
the treatment system (8). Optimization maximizes the efficient use of chemical,
minimizes additional solids production, and produces soluble phosphorus residuals of
less than 0.5 mg/1. Hais, et al. (27), indicated, using step aeration with alum addition,
that above pH 6.6 a rapid increase in the effluent phosphorus residual occurred with
increasing pH. However, operation with a mixed liquor pH of less than 6.2 caused an
upset in the activated sludge process. It was concluded that optimum performance (0.6
mg/1 effluent phosphorus residual) was achieved by maintaining a final stage pH
between 6.3 and 6.6 at an A1:P weight ratio of 2:1. Mulbarger (8) showed the
capability of pH adjustment with a denitrification system and demonstrated the
advantages of split chemical treatment in a multi-sludge system. Results from the pH
adjusted system are shown in Figure 7-8. Long, et al. (9), found a natural pH
adjustment with high doses of alum in a poorly buffered wastewater and report
phosphorus residuals below 0.5 mg/1 at A1:P weight ratios of from 2:1 to 3:1.
Organism activity and type (9) as well as the completeness of a chemical reaction
affect phosphorus removal and are influenced by the temperature. No temperature
effect has been reported, however, for the rate of chemical reaction (13) operating
between 10 and 20° C.
Mulbarger (8, 16) notes that the presence or absence of nonsettleable effluent
suspended solids from a combined chemical-biological system is more dependent upon
the ratio of net volatile solids produced to aluminum added than upon pH, exceeding
an aluminum to phosphorus ratio, or exceeding a given aluminum dosage. It is
recommended that this ratio not drop below 3 to 5. It is believed that the biological
volatile solids provide exchange and/or sorption sites of limited capacity for the
precipitated aluminum-hydroxy-phosphate complex, and that these natural polymers aid
in clarification. Thus, the more biological solids produced from the system the greater
the aluminum dosage that can be used without effluent suspended solids problems.
When designing the phosphorus removal system, the designer should be cognizant of
the phosphorus levels in the effluent suspended solids. There was a nominal 1.0 mg/1
of phosphorus in the effluent suspended solids from the high rate system at Manassas
(8). After pH adjustment and addition of 0.3 to 0.4 mg/1 of anionic polymer (Nalco
676), effluent suspended solids were reduced to 2 mg/1. Consequently, the effluent
total phosphorus was also reduced. Polymer performance was also reported to be
influenced by the pH (8). If total phosphorus residuals of less than 0.5 mg/1 as P are
required, a multimedia filtration system is recommended. Dual and tri-media filtration
both provide effective tertiary solids separation. Hais, et al. (27) showed that tri-media
filtration consistently removed between 5 and 10% more suspended material than did
the dual-media filter. Filter runs between 24 and 32 hours were regularly maintained.
Tri-media filtration performance was reported to be superior to dual-media filtration
at Manassas (8).
7 - 17
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1.0
2.0 3.0
ALUMINUM ADDED
4.0
5.0
Ib
CHEMICALLY REMOVABLE PHOSPHORUS' Ib
FIGURE 7-8 EFFECT OF ALUMINUM ADDITION WITH pH
ADJUSTMENT ON TOTAL SOLUBLE PHOSPHORUS RESIDUAL
7.3.2 Costs
Power costs for combined chemical-biological treatment are negligible. Maintenance
costs will increase somewhat due to the additional equipment. Capital expenditures for
storage facilities, instrumentation, and equipment will be very low at larger facilities
unless multi-media filtration is incorporated. Chemical costs will vary as a function of
the incoming phosphorus concentration and the required phosphorus residual.
Additional expenditures for pH control (alkalinity dependent) and solids control
(polymer additions) may be necessary. See Chapter 10 for further information.
7.4 Process Design Examples
Four design example problems were selected for an arbitrary wastewater. Each
illustrates specific points in combined chemical-biological wastewater treatment with
alum additions for phosphorus removal. Results from the Manassas work are used for
some of the calculations.
Assumptions:
a. Net biological solids production = 0.28 Ib of volatile solids produced per Ib
of COD removed; the volatile solids contain 2% phosphorus; and the total
solids average 80% volatile.
b. The biological system is to be designed with a minimum hydraulic contact
time of 1 hour at the maximum 4 hour hydraulic peak. The ratio of
7-18
-------
maximum flow rate to average flow rate = 1.5. A cell retention time (CRT)
of two days is selected.
c. Primary Effluent Characteristics:
COD = 200 mg/1
P = 10 mg/1
Alkalinity = 150 mg/1 as CaCO3
Find:
1. Effluent COD, BOD5 and phosphorus
2. Total solids production
3. Alkalinity loss
4. Dissolved solids addition
5. Size of final clarifier, return and waste sludge pumping systems.
6. Chemical costs with the following delivered prices:
Alum = $0.32/Ib A13+
Polymer = $1.50/lb polymer
Acid = $0.02/lb H2SO4
Lime = $0.01/lb Ca(OH)2
Example No. 1: High rate system for most efficient chemical addition with pH
above 7 and no optimization by pH adjustment.
1. Effluent COD, BOD5, and Phosphorus
(See Table 7-4. Note: Soluble organics from a biological system vary as a
function of the wastewater and dilution of the refractory fraction. Manassas
values are used for illustration only.)
Effluent soluble COD = ^227152 <41'33) + 33
5 + 33 = 38 mg/1
Assume effluent suspended solids = 20 mg/1
Total effluent COD = 20 (0.58) + 38
= 12 + 38
= 50 mg/1
7-19
-------
Effluent BOD5 = 38 (0.28) + 20 (0.23)
= 10.6 + 4.6 [Using BOD5 ratios
from Table
Background phosphorus removal per Ib COD removed
= (biological solids production ratio) (phosphorus content)
= (0.28) (0.02)
= 0.0056 Ib phosphorus removed
Ib COD removed
[The background phosphorus removal rate at Manassas was found to be approximately
0.01 Ib P Removed
Ib COD removed ( } J
Background phosphorus removal expressed as concentration
= (COD removal) (phosphorus removal per Ib COD removed)
= (200-38) (0.0056)
= 0.9 mg/1
Chemically removable phosphorus = 10.0 - 0.9 = 9.1 mg/1
Use Al /chemically removable phosphorus dose = 1.25 Ib/lb
From Figure 7-6 find 0.6 Ib of phosphorus removed per Ib of aluminum added at the
dosing ratio of 1.25
Aluminum dose = 1.25 (9.1) = 11.3 mg/1 A13+
Chemical phosphorus removal = 0.6 (11.3) = 6.8 mg/1 P
Soluble phosphorus residual = 9.1 - 6.8 = 2.3 mg/1
Total effluent phosphorus = 2.3 + 20 (0.05) [P/SS ratio from
= 3.3 mg/1 P Table 7-4]
The preceding calculations are useful from the standpoint that they show the pollutant
fractions in both the soluble and solid phases. Often a remarkable treatment im-
provement can result from effluent solids control and minimization.
2. Alkalinity Loss
Maximum possible alkalinity loss due to aluminum hydrolysis
= 5.4 (5.5) = 30 mg/1 as CaCO3
which gives ~ = 2.7 Ib CaCO3 loss/lb A13+
7-20
-------
At an aluminum added to chemically removable phosphorus ratio of 1.16, an average
alkalinity loss of 2.3 Ib CaCOg/lb AP+ was measured at Manassas (8). Like solids
production values, alkalinity losses are a function of the proximity of aluminum dose to
stoichiometric phosphorus requirements and the reaction pH.
3. Dissolved Solids Addition
9 , mg SO42"
SO/' = (11.3 mg A1J+/1) (5.4 £- ) = 61 mg/1
* mg A13+
Although dissolved solids are added to the system, the designer should remember that
other soluble wastewater components (i.e., organics, phosphorus, alkalinity, etc.) are
being removed. Prior to the addition of sulfuric acid for pH adjustment at Manassas, no
increase in dissolved solids was measured on a primary to final effluent basis because of a
compensating loss of soluble wastewater components (8).
4. Final Clarifier and Pumping System
Organic solids = 57 mg/1
Inorganic solids = 43 mg/1
Total = 100 mg/1
Solids lost in effluent = 20 mg/1
Solids accumulation = 80 mg/1
CRT = 2 days, and aeration tank detention at average flow
= 1.5 (1) = 1.5 hours
(24)(2)
MLSS in aerator = ' (80) = 2,560 mg/1
Since this is the final clarifier a conservative overflow rate should be utilized. From
Figure 7-9 the minimum settling rate of 1,150 gpd/ft2 was observed at a MLSS of 2,600
mg/1. To provide for peak flow, the clarifier should be designed for an overflow rate of
^
770 gpd/ft , or less. It should have a 12 ft effective water depth, and might be provided
with a surface skimmer.
Maximum SVI at 2,600 mg/1 MLSS = 82 ml/gm (Figure 7-9)
Approximate return sludge concentration at worst condition
106 106
= 12,200 mg/1
SVI 82
Solids balance equation where R is sludge return rate
+ R) (MLSS) <_ (R) (return sludge cone.)
7-21
-------
E
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75
50
25
1500
1000
500
MAXIMUM
MINIMUM
MAXIMUM
MINIMUM
AVERAGE RANGE OF OPERATING
SURFACE OVERFLOW RATE: 680-13W gpd/ft2
1000 2000 3000 4000 5000
MIXED LIQUOR SUSPENDED SOLIDS, mg/1
6000
FIGURE 7-9 PHYSICAL CHARACTERISTICS OF HIGH RATE
ACTIVATED SLUDGE WITH MINERAL ADDITION (8,16)
7-22
-------
or considering just the case of equality,
2,600
' gave 12,200-2,600
R = 0.4 Qaye
(Actually, effective design would indicate a ±50% capability; i.e.,
0.2 to 0.6 Qaye capability)
Solids to be wasted per day = 80 mg/1 at a minimum sludge solids concentration = 12,200
mg/1. Assume continuous wastage with separate pump.
80
••* T X-* *-• V _ f\ f\f\ /"C f~\
w ~ ^ave 19 ?nn ~ u-uuo:> ^ave
(Again, effective design would indicate a ±50% capability; i.e., 0.003
to 0.01 Qave capability).
5. Alum Cost
= (11.3)(0.32)(8.34)(10~3) = 3.1 #1.000 gal.
Additional Comments
No mention has been made of the effect of sidestream phosphorus contributions. With
conventional solids handling and treatment processes, gravity and flotation thickening,
filtration and centrifugation, aerobic and anaerobic digestion, and incineration,
phosphorus recycle should not be a problem and any that occurs will largely be
associated with the solids fraction of the minor process streams. However, processes that
cause biological solids liquefaction and hydrolysis, such as heat treatment processes, will
result in a substantial recycle of soluble phosphorus with a corresponding increase of
metal ion requirements.
Example No. 2: Same system as Example No. 1 but with sulfuric acid addition
for pH adjustment (residual alkalinity = 30 mg/1 as CaCOo) and polymer addition
for solids control.
Detailed calculations are not shown for the problem since the techniques are essentially
the same as in Example No. 1. Results are summarized, along with results from the other
examples, in Table 7-5. Highlights from this problem are as follows:
1. COD, BODj, and suspended phosphorus concentrations in the effluent are
substantially reduced by an assumed polymer dose of 0.3 mg/1. The
chemically removed phosphorus was calculated for this case by multiplying the
0.6 Ib P removed/lb A13+ added, from Figure 7-6, by 1.33. The 1.33 was
obtained from Figure 7-4, assuming an adjusted pH of about 6. The result was
a phosphorus residual of 0.1 mg/1 with an aluminum dose of 11.3 mg/1. But
7-23
-------
Figure 7-8 showed from experimental data a nominal 0.5 mg/1 residual
phosphorus at this A1:P ratio and, therefore, the conservative residual of 0.5
mg/1 soluble phosphorus was chosen.
Table 7-5
SUMMARY OF DESIGN CALCULATIONS
Item Example 1 Example 2 Example 3 Example 4
pH Adjustment No Yes Yes Yes
Polymer Addition No Yes Yes Yes
Split Treatment* No No No Yes
Mixed Media Filtration No No Yes Yes
Total Aluminum Dose, mg/1 11.3 11.3 27.3 18.2
Effluent Quality, mg/1
Suspended Solids 20 2 0 0
COD 50 39 38
BOD5 15 11 11
Total Phosphorus 3.3 0.6 0.2 0.2
Total Soluble Phosphorus 2.3 0.5 0.2 0.2
Solids Production, mg/1
Biological 57 57 57 57+
Chemical 43 45 92 66
Alkalinity Loss
Due to A13+ Addition 30 21 109 59
mg/1 as CaCO3
o
804 Increase Due
to A13+ Addition, mg/1 61 61 147 98
Alum Cost, tl 1,000 gal. 3.1 3.1 7.4 4.9
Polymer Cost, 4 / 1,000 gal. 0 0.3 0.3 0.3
Acid Cost, 41 1,000 gal. 0 1.7 0 1.0
Total Cost, tf / 1,000 gal. 3.1 5.1 7.7 6.2
Clarifier Overflow Rate, gpd/ft2 770 670 360
Return Sludge Rate 0.4Q 0.6Q 0.9Q
Waste Sludge Rate 0.0065Q 0.008Q 0.012Q
*Split treatment here incorporates mineral addition as in Example No. 1 without pH
adjustment, followed by mineral addition with pH adjustment at a later point.
7-24
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2. The alkalinity loss due to hydrolysis was also reduced due to increased
formation. Using 96% strength sulfuric acid and assuming a residual alkalinity
of 30 mg/1 as CaCOo, a calculated sulfuric acid requirement of 101 mg/1 was
found. The sequence of chemical addition is acid, alum, and polymer, with
about 5 minutes of mixing at peak flows between each application.
o
3. Dissolved solids additions totaled 156 mg/1 SO4 ; 61 mg/1 due to the alum and
95 mg/1 due to the acid.
4. Because of improved solids capture, the operating MLSS in the aerator
increased to 3,200 mg/1 necessitating a lower surface overflow rate in the
clarifier, and greater return and waste sludge handling capability.
5. Chemical costs increased to 5.1^/1,000 gal. with benefits in phosphorus,
suspended solids, BOD^ and COD removals.
Example 3: Assuming mixed media filtration provides complete solids control,
what is the change in system calculations if a residual effluent phosphorus of 0.2
mg/1 as P is to be provided?
Highlights from this problem are as follows:
1. Essentially no differences from Example No. 2 are found in effluent COD and
BOD. However, the mixed media filter does provide positive effluent solids
control in the event of system upset. An A1:P weight ratio of 3 (Figure 7-8)
was used.
2. The addition of alum caused an alkalinity consumption of 109 mg/1, which
would deplete the background alkalinity of the wastewater to 41 mg/1 as CaCOo.
If a substantially higher alum dosage had been required, all of the
alkalinity might have been consumed, and addition of lime might have
been necessary.
<-}
3. Dissolved solids addition consisted of 147 mg/1 804 from the alum.
4. Additional solids productions resulted in an operating solids level of 4,800 mg/1
in the aerator, necessitating a lower surface overflow rate and higher return
and waste sludge capabilities than in earlier examples.
5. Total chemical treatment costs are raised to 7.7^/1,000 gal.
Example No. 4: Split treatment (chemical addition at more than one point) is
utilized for a required effluent phosphorus concentration of 0.2 mg/1. Assume the
first point of addition is as in Example No. 1. Further mineral addition will be
made to the effluent in Example No. 1 either before filtration or during a second
stage of biological treatment.
7-25
-------
1 . Changes in effluent COD and BOD are not calculated since this is dependent
upon whether split treatment is practiced with just filtration, or whether
another biological system is added. In a real situation it would be better to
reduce the aluminum dose below that in Example No. 1 in the first stage
system for more efficient chemical use. This was not done here for design
simplicity. Another possibility is addition of part of the chemical dose in the
primary settler.
The Al dose of 6.9 mg/1 for the second addition was arrived at using Figure 7-8
for a pH adjusted system.
2. The alkalinity depletion due to aluminum addition was calculated as 59 mg/1
as CaCOo. This would leave a residual alkalinity of 91 mg/1. Sixty mg/1 as
FUSO^ was added to reduce the alkalinity to 30 mg/1. Again, in a real
situation acid might not have to be added to this wastewater.
^
3. Dissolved solids additions were calculated as 98 mg/1 SO4 due to alum
9-
addition and 59 mg/1 SO^ due to acid addition.
4. Clarifiers were not sized in this example. If two biological processes were used
in series, the final clarifier of the first process could be designed with an
overflow rate from the middle' of the performance curve in Figure 7-9.
5. The cost of acid is l.Oi / 1,000 gal. Total chemical costs are 6.2tf / 1,000 gal.
Comment: This final example, in comparison to example No. 3, shows the benefit
derived from split chemical treatment for low phosphorus residuals.
7.5 References
1. Guggenheim Process Described by: Kiker, J. E., "Waste Disposal for Dairy Plants",
Sanitarian (Los Angeles) 16, p 11 to 17 (1953). Florida Engng. Exp. Station 7,
Leaflet Series No. 51 (1953). Moore, R. B., "Biochemical Treatment for Anderson,
Ind.", Wat. Works and Sew., 85:11 (1938).
2. Thomas, E. A., "Phosphat-Elimination in der BelebtechlemannlagevonMannedorf und
Phosphat-Fixation in See-und Klarschlamm", Vierteljahrsschrift der Naturforschenden
Gesellschaft in Zurich, 110, Schlussheft, S., p 419 (1965).
3. Wirts, J. J., "Communications", Buckeye Bulletin (Summer, 1966).
4. Tenny, M. W., and Stumm, W. J., "Chemical Flocculation of Mirco-organisms in
Biological Waste Treatment", JWPCF, 37:10, p 1370 (1965).
5. Barth, E. F., and Ettinger, M. B., "Mineral Controlled Phosphorus Removal in the
Activated Sludge Process", JWPCF, 39:8, p 1361 (1967).
6. Eberhardt, W. A., and Nesbitt, J. B., "Chemical Precipitation of Phosphorus in a High
Rate Activated Sludge System", JWPCF, 40:7, p 1239 (1968).
7-26
-------
7. Barth, E. F., Brenner, R. C., and Lewis, R. F., "Chemical-Biological Control of
Nitrogen and Phosphorus in Wastewater Effluent", JWPCF, 40:12, p 2040 (1968).
8. Mulbarger, M. C., "The Three Sludge System for Nitrogen and Phosphorus Removal",
To be presented at the 44th Annual Conference of the Water Pollution Control
Federation, San Francisco, California (October, 1971).
9. Long, D. A., Nesbitt, J. B., and Kountz, R. R., "Soluble Phosphate Removal in the
Activated Sludge Process — A Two Year Plant Scale Study", Presented at the 26th
Annual Purdue Industrial Waste Conference, Purdue University, LaFayette, Indiana
(May, 1971).
10. Connell, C. H., and Frey, S. M., Private Communications (January, 1971).
11. Humenick, M. J.,and Kaufman, W. J., "An Integrated Biological-Chemical Process for
Municipal Wastewater Treatment", Presented at the 5th International Water Pollution
Conference, San Francisco, California (July, 1970).
12. Schmidt, F., and Ewing, L., "Phosphate Removal System for Small Activated Sludge
Plants", Presented at the Pennsylvania Water Pollution Control Association Meeting,
State College, Pennsylvania (August, 1970).
13. Recht, H. L., and Ghassemi, M., "Kinetics and Mechanism of Precipitation and
Nature of the Precipitate Obtained in Phosphate Removal from Wastewater Using
Aluminum (III) and Iron (III) Salts", Water Pollution Control Research Series, 17010
EKI, Contract 14-12-158, USDI, FWQA (April, 1970).
14. Rickert, D. A., and Hunter, J. V., "Effects of Aeration Time on Soluble Organics
During Activated Sludge Treatment", JWPCF, 43:1, p 134 (1971).
15. Gulp, R. L., and Gulp, G. L., Advanced Wastewater Treatment, Van Nostrand
Reinhold Company, New York (1971).
16. Mulbarger, M. C., and Shifflett, D. G., "Combined Biological and Chemical
Treatment for Phosphorus Removal", Chem. Eng. Prog. Symp. Series, 67:107, p 107
(1970).
17. Sawyer, C. N., Chemistry for Sanitary Engineers, McGraw-Hill, New York (1960).
18. McKee, J. E., and Wolf, H. W., Water Quality Criteria, Publication No. 3-A,
California Water Quality Board, Sacramento, California (1963).
19. Cherry, A. L., and Schuessler, R. G., "Private Company Improves Municipal Waste
Facility", Wat. and Wastes Eng. 8:3, p 32 (1971).
20. Water Quality and Treatment, 2nd Ed., American Water Works Association, New
York (1951).
21. Ockershausen, R. W., Harriger, R. D., and Zuern, H. E., "Phosphorus Removal Tests
with Alum" for Buffalo Sewer Authority, Buffalo, New York, by Technical Service
7-27
-------
Department, Allied Chemical Corporation, Industrial Chemicals Division, Morristown,
New Jersey (1971).
22. Fair, G. M., and Geyer, J. C., Water Supply and Wastewater Disposal, Wiley and
Sons, New York (1961).
23. Detroit Metro Water Department, "Development of Phosphate Removal Processes",
Detroit, Michigan (Program 17010 FAH Grant WPRD 51-01-67 Advanced Waste
Treatment Laboratory, Cincinnati, Ohio), Prep ub lie at ion Copy (July, 1970).
24. Boggia, C., and Herriman, G. L., "Pilot Plant Operation at Warren, Michigan",
Proceedings of the 43rd Annual Conference of the Michigan Pollution Control
Association (1968).
25. Leary, R. D., Ernest, L. A., Powell, R. S., and Manthe, R. M., "Phosphorus Removal
with Pickle Liquor in a 115 MGD Activated Sludge Plant", Sewerage Commission of
the City of Milwaukee, Wisconsin, Grant No. 11010FLQ, Water Quality Office,
Environmental Protection Agency, Advanced Waste Treatment Research Laboratory,
Cincinnati, Ohio, Prepublication Copy, (1971).
26. Grigoropoulos, S. G., Vedder, R. C., and Max, D. W., "Fate of Aluminum-Precipitated
Phosphorus in Activated Sludge and Anaerobic Digestion", Presented at the 43rd
Annual Conference of the Water Pollution Control Federation, Boston, Massachusetts
(1970).
27. Hais, A. B., Stamberg, J. B., and Bishop, D. F., "Alum Addition to Activated Sludge
with Tertiary Solids Removal", Presented at the 68th National Meeting of the AIChE,
Houston, Texas (March 1971).
28. Dean, R. B., "Sludge Handling", Presented at the Advanced Waste Treatment and
Water Reuse Symposium, 4th Session, Sponsored by the Environmental Protection
Agency, Dallas, Texas (Jan. 12-14, 1971).
29. Smith, J. E. Farrell, J. B., and Dean, R. B., Unpublished Data from the Advanced
Waste Treatment Laboratory, Office of Research and Monitoring, Environmental
Protection Agency, Cincinnati, Ohio (Jan., 1971)
30. Dotson, G. K., Unpublished Data from the Advanced Waste Treatment Laboratory,
Office of Research and Monitoring, Environmental Protection Agency, Cincinnati,
Ohio (July, 1971).
7-28
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Chapter 8
PHOSPHORUS REMOVAL BY LIME TREATMENT
OF SECONDARY EFFLUENT
8.1 Description of Process
8.1.1 Theory
Lime treatment of wastewater is essentially the same process as the familiar lime
softening of drinking water supplies. The objectives, however, are quite different. While
softening may occur, the primary objective is to remove phosphorus by precipitation as
hydroxyapatite. This reaction was described in Chapter 3.
During phosphorus precipitation other important reactions occur. The reaction of lime
with alkalinity, that results in calcium removal when carrying out softening, not only
takes place when treating wastewater, but may have a very important effect on the
general efficiency of the process. This reaction can be considered to take place in the
two following ways:
Ca(HC03)2 + Ca(OH)2 * 2 CaCOg + 2 H2O
NaHCO3 + Ca(OH)2 > CaCO? + NaOH + H2O
The first equation is that for softening. Some wastewaters do not contain enough
calcium, however, for that equation to be satisfied. Calcium carbonate precipitation
may still occur, but by the second reaction. The reactions to form CaCO^ are
important for two reasons: the lime consumption determines to a considerable extent
the lime dose required for operating the process, and the resulting CaCO^ acts as a
weighting agent to aid in settling of sludge.
Another reaction that may be important is precipitation of Mg(OH)2 as follows:
Mg2+ + 2 OH" -» Mg(QH)2
This reaction does not approach completion until the pH is raised to 11. Magnesium
hydroxide is a gelatinous precipitate which aids colloid removal, but which hinders
sludge thickening and dewatering. Where the pH must be raised to 11 or above to
meet a phosphorus removal requirement or some other treatment objective, Mg(OH>2
formation must be considered.
In the two-stage lime treatment process which will be described below, it is necessary
to include recarbonation after the first stage to reduce the pH and precipitate the
excess lime as CaCOo in the second stage. The following reaction occurs:
Ca2+ + CO2 + 2 OH' * CaCO3 + H2O
Carbon dioxide may also be used to lower pH after lime treatment. The important
^
reaction in this case would be the conversion of Co^ to HCO3".
- 1
-------
A final reaction that is of major concern, where lime is to be recovered by sludge
recalculation, is as follows:
CaCO3 A* CaO + CO2
The CaO produced would then be slaked to form Ca(OH>2 before use.
8.1.2 Treatment Systems
Two lime treatment systems may be used with wastewater, single-stage and two-stage.
Figure 8-1 shows a single-stage system. In single-stage treatment lime is mixed with
feed water to raise the pH to a desired value. Although the pH will depend upon the
required phosphorus removal, it is likely to be substantially less than 11, and may be
less than 10. Precipitation of phosphate and other materials, as indicated by reactions
discussed above, takes place. Time is allowed in the appropriate equipment for
precipitate particles to flocculate to sufficient size for good settling. The clarified water
from the settler may be discharged directly or may be filtered to improve solids
removal. Adjustment of pH with CO2 may be necessary before discharge and will
almost certainly be required before filtration to prevent post-precipitation of CaCO3
from the unstable water. The settled lime sludge may be disposed of as landfill or
may be recalcined for recovery of lime. In the latter case, the sludge is thickened,
dewatered by centrifuge or vacuum filter, and calcined. The calcined product is then
slaked and reused. To avoid buildup of inerts in the lime, some of the sludge or
recalcined lime must be wasted or the inerts must be separated from the sludge before
calcination.
Two-stage treatment is somewhat more complicated than single-stage treatment. In
typical two-stage treatment, shown in Figure 8-2, enough lime is added to the water in
the first stage to raise the pH above 11. Precipitation of hydroxyapatite, CaCO3, and
Mg(OH>2 occurs. Consideration of the solubility product for CaCOo and the
equilibrium between CO^ and HCO3~ shows that the minimum solubility for Ca
occurs at a pH of about 10. At a pH of 11 or above there is a considerable
0-4-
concentration of Ca present in the water. In two-stage treatment CO2 is added after
the first-stage settler to bring the pH down to about 10 where CaCO^ precipitation
results. The CaCO-j is settled out and the clarified water is either discharged or sent
to filtration. As in the case of single-stage treatment, pH reduction may be necessary
before discharge and probably would be required before filtration.
8.2 Typical Performance Data
A number of full-scale and pilot-scale tertiary lime treatment plants are in operation
and more full-scale plants are beginning operation. Only one plant, that at South Lake
Tahoe, California, has operated for a significant period with recalcination of sludge for
lime recovery. Most aspects of this 7.5 mgd plant have been discussed in detail by
Gulp and Gulp (1). Although the data on handling of wastewater sludges for
recalcination and the recalcination process itself are limited mainly to this plant,
8-2
-------
CO
CO
I—
LU
-------
CO
>-
to
-------
performance data on other parts of a lime treatment system are available from more
than one location. Extensive data have been reported from pilot plants located in
Washington, D. C., (2) and Lebanon, Ohio (3).
The lime treatment system at Tahoe is a two-stage clarification system which is
followed by pressure filters of the multimedia type. These contain 3 ft. of a mixture
of coal, sand, and garnet. The water passes through two filters in series. The first stage
clarifier is operated at a pH of 11. To reach that pH requires about 300 mg/1 CaO. A
small amount of polymer is used to improve flocculation. The pH is reduced to 9.6 by
recarbonation before the second-stage settler. Before filtration, the pH is further
reduced to 7.5. A small amount of alum (1 to 20 mg/1) or a combination of alum
and polymer is used as filter aid. Filter run length has varied between 4 and 60 hours.
Phosphorus removal at this plant always has been good and has improved as operating
experience has increased. By returning plant waste streams containing precipitated
phosphorus to the first stage flocculator, it is now possible to obtain routinely an
effluent with less than 0.1 mg/1 P. Before the filters the phosphorus concentration is
about 0.4 mg/1.
In addition to phosphorus removal, there is significant removal of organic materials.
During a period of intensive study of the first stage of clarification (unpublished data)
77% removal of BOD and 61% removal of COD were obtained. Although there was
not a large removal of suspended solids, there was a significant change in the character
of the solids from organic to largely inorganic. Further removal of organic materials
and suspended solids occurs over the remainder of the system. Gulp and Gulp report
typical filter effluents with a BOD of 3 mg/1, COD of 25 mg/1, and turbidity of 0.3
JTU.
Sludges from the first stage settler and the settler following recarbonation are sent to a
gravity thickener. Solids concentration increases during thickening from about 1% to
from 8 to 20%. Thickened sludge is then centrifuged to form a cake with from 30%
to more than 40% solids. The cake is next calcined in a multiple hearth furnace and
the recovered lime slaked for reuse. Initially it was planned to operate the centrifuge
for high solids recovery in the cake. This necessitates discarding a significant amount
of the recovered lime to prevent buildup of precipitated phosphate and other inerts.
Since operation began, it has been found that much of the phosphate and some other
inert materials can be separated from the CaCOg in the sludge by operating with a
rather high solids concentration in the centrate. Approximately 90% of the phosphorus
can be removed in the centrate, for example, when 25% of the solids entering the
centrifuge are allowed to remain in that stream. This classification procedure results in
a loss of about 15% of the recoverable lime. A greater degree of utilization of
recalcined lime compensates for the loss, however, and the load on the calcining
furnace is reduced. Over three years of operation, mostly without classification in the
centrifuge, the average concentration of CaO in the recalcined product was 66%.
Recalcined lime made up 72% of the total used at the plant.
- 5
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Considerable operating data have been obtained by O'Farrell and Bishop (2) on a
two-stage system operated at the EPAWQO-DC Pilot Plant in Washington, D. C. This
50,000 gpd plant treated secondary effluent from a pilot activated sludge system at
the municipal treatment plant. The first stage pH was maintained at about 11.7, using
a lime dose of about 400 mg/1 as CaO, and the pH after recarbonation was
maintained at about 10.3. Ferric iron was added at a rate of 5 mg/1 in the second
stage to improve flocculation. Water from the second-stage settler was filtered without
further pH adjustment through gravity-flow dual-media filters consisting of anthracite
coal over sand. The average filter run length was more than 50 hours.
Phosphorus removal for the system was similar to that obtained at Tahoe; 0.09 mg/1
as P was the average concentration remaining. Phosphorus concentration before
filtration averaged 0.13 mg/1.
In addition there was significant organic and suspended solids removal. The average
BOD was reduced from 15 mg/1 to 2.1 mg/1 before and 1.5 mg/1 after filtration.
Suspended solids were reduced from 33 mg/1 to 17 mg/1 before filtration and 3.8
mg/1 after filtration.
Sludge from the second stage settler was returned to the first stage and all sludge
removed from the system was from the settler of the first stage. This sludge contained
about 5% solids and constituted about 1.5% of the feed volume. Although there was
sludge thickening and calcining equipment available, only preliminary study was made
of lime reuse. Limited results showed improved thickening of the sludge when
recalcined lime was recycled to the system.
A 75 gpm single-stage pilot system was operated at the Lebanon, Ohio Sewage
Treatment Plant. Feed water was activated sludge effluent. This system, consisting of a
flocculator-clarifier followed by dual-media filters of anthracite coal and sand, is
described by Berg, et al. (3). Sludge was gravity thickened and discharged to sand
drying beds. The system has been operated over a pH range from about 9 to 11 with
most data taken at a pH of 9.5. Before filtration the pH was reduced to about 8.8
with sulfuric acid to prevent precipitation on the filter media. Filter run length
averaged about 90 hours.
Effluent phosphorus concentrations for the system are shown in Figure 8-3. The effect
of pH is clearly indicated. Since the average phosphorus concentration of the
secondary effluent was about 10 mg/1 as P, approximately 95% removal was obtained
at a pH as low as 9.5. Before filtration the average phosphorus concentration was 0.75
mg/1 at a pH of 9.5. Removal improved only slightly at higher pH because of the
presence of precipitated phosphorus in the effluent.
Average suspended solids concentration of the secondary effluent was 43.5 mg/1. This
was reduced to 16.5 mg/1 in the clarifier. Much of the suspended matter in the clarifier
effluent consisted of inorganic precipitates. After filtration the water was of high
clarity. During a six month period when the biological treatment plant operated well,
turbidity averaged 0.2 JTU. Even when the suspended solids load from the activated
8-6
-------
2.0
1.5
en
E
1.0
o
o
DC
O
3:
Q-
o
a:
a.
0.5
8.5
9.0
9.5 10.0 10.5
CLARIF IER pH
11.0
11.5
FIGURE 8-3 EFFECT OF pH
ON PHOSPHORUS CONCENTRATION
OF EFFLUENT FROM FILTERS
FOLLOWING LIME CLARIFIER
- 7
-------
sludge plant was high, the lime treatment system produced water with a daily average
turbidity not exceeding 2.5 JTU.
During a two month period in which extensive measurement of organic removal was
made, total organic carbon was reduced by 63% over the clarifier and 68% over the
clarifier and filters. Organic carbon measurements taken at other times indicated
removals of from 55 to 74%. Occasional BOD and COD samples indicated removals of
86% and 62%.
Sludge was usually removed from the settler at 1.5 gpm giving a sludge concentration
of 2.8% solids. This was thickened to about 10% before being pumped to drying beds.
The sludge dewatered quickly to about 50% solids. Equipment was not available for
recalcining the sludge.
The lime requirement is an important consideration in lime treatment. This can vary
over a wide range depending on operating pH and water composition. From reactions
given earlier it was seen that alkalinity has an important effect on the lime dose.
Buzzell and Sawyer (4) have shown, for example, that for several wastewaters the lime
dose required to reach a pH of 11 correlated approximately with alkalinity.
Examination of the earlier equations would show also that calcium hardness can affect
lime dose. One part by weight of CaO can react with from 0.89 to 1.79 parts of
bicarbonate alkalinity expressed as CaCOo, the lower value applying to very soft waters
and the higher value to very hard waters. In addition to the reaction of lime with
hardness, other competing reactions occur in lime treatment of wastewater. Also, there
may be incomplete reaction of the lime. All of these complications make calculation
of lime dose difficult. The result is that, at present, determination of lime dose is
largely empirical. Approximate values have already been given for the plants at Tahoe
and Washington, D.C. Some approximate values for the Lebanon work are shown in
Table 8-1. It appears from the data already available that the lime dose will usually be
in the range of 300 to 400 mg/1 as CaO for two-stage treatment, and from 150 to
200 mg/1 where single-stage treatment is satisfactory.
Table 8-1
Feed Water
Alkalinity
(mg/1 as CaCO3)
300
300
400
400
LIME REQUIREMENTS
Clarifier pH
Approximate Lime
Dose
9.5
10.5
9.5
10.5
(mg/1 of CaO)
185
270
230
380
8 -8
-------
8.3 Criteria for Selection of Process
Lime treatment of secondary effluent represents a significantly higher capital cost and
somewhat higher operating cost for phosphorus removal than mineral addition to a
conventional treatment plant. Lime treatment has, however, several advantages over
mineral addition. It can be considered more dependable since it adds additional
flocculation and sedimentation steps to the system. Upsets in the conventional plant,
which would reduce the efficiency of phosphorus removal by mineral addition, would
have less effect on tertiary lime treatment. Because tertiary lime treatment is separate
from the conventional treatment plant, it adds flexibility to operation of the system.
For very high degrees of phosphorus removal, lime treatment would be the method of
choice, not only because of the inherent greater dependability mentioned above, but
because of the ability to produce an effluent slightly lower in phosphorus content.
Lime treatment decreases the total dissolved solids content of the water by removal of
hardness and alkalinity while mineral addition adds to the total dissolved solids. In the
case of alum addition, for example, 5.3 parts by weight of sulfate are added for each
part of aluminum. Lime treatment has the capability of removing turbidity to very low
levels. Where there are plans for recreational reuse or certain industrial reuses of the
treated water, low turbidity along with low phosphorus content of the lime treatment
effluent is very desirable. Lime treatment offers the opportunity for recovery of the
treatment chemical. At the present time there are no acceptable methods for recovery
of aluminum or iron salts.
The choice of single-stage or two-stage lime treatment depends partly upon the degree
of phosphorus removal required, but more importantly, on the alkalinity of the water.
Unless a high treatment pH is used, waters with low alkalinity, in the range of 150
mg/1 as CaCOo or less, form a poorly settleable floe because of the low fraction of
dense CaCOo. A pH above 10 is needed even to obtain measurable CaCO^ production.
A pH of 11 or above is then needed to precipitate Mg(OH)2 which aids in settling of
fine particles. Since neither discharge nor reuse of the high-pH water is likely to be
acceptable, addition of a second treatment stage after recarbonation to a pH of about
10 becomes generally necessary. It would be possible to recarbonate to a much lower
pH and avoid the second treatment stage, but this would result in a high calcium
effluent and would eliminate the chance to produce high CaCOo sludge in situations
J
where lime recovery was contemplated. Although most of the phosphorus removal
occurs in the high pH first stage, in accordance with the pH dependence of
phosphorus solubility as shown earlier by Figure 8-3, some removal does also occur in
the second stage. Another important reason for including the second stage is to assure
better control of clarification. With low alkalinity waters there is sometimes difficulty
in settling the sludge in the first stage even at high pH. The second stage settler with
its high CaCOo sludge, prevents solids carryover when the first stage settler does not
operate properly.
With high alkalinity waters, a well settling floe is formed at pH values as low as 9.5.
There is no need for two-stage treatment unless the required degree of phosphorus
removal necessitates a high pH. In addition to obtaining a small degree of phosphorus
8-9
-------
removal, the second stage would be used in these cases to lower calcium content in
the effluent and to obtain high CaCOg sludge for lime recovery. There is not yet
available operating experience at enough plants to state positively the alkalinity at
which single-stage treatment will perform satisfactorily, from the standpoint of floe
settleability. At present, experience indicates that at an alkalinity of 150 mg/1 as
CaCOg, single-stage treatment probably cannot be used. Between 150 mg/1 and 200
mg/1 settleability will probably depend on the amount of organic floe present. Above
200 mg/1, settleability is likely to be satisfactory.
8.4 Description of and Criteria for Choice of Equipment
8.4.1 Single-Stage System and First Stage
of a Two-Stage System
Diagrams for single and two-stage lime treatment systems are shown in Figures 8-1 and
8-2. The first part of each system is made up of clarification equipment essentially the
same as that found in water treatment plants. It includes a chemical feeding system
(see Chapter 10), a rapid mix tank, a flocculator, a settler, and a means for sludge
removal. Mixing, flocculation, and settling may be carried out in separate vessels or
tanks, or they may be combined in one integrated unit. The system consisting of
separate tanks has been used for many years in water treatment and was the system
adopted for the first stage of the Tahoe lime treatment plant. The integrated type of
equipment is sometimes referred to as an upflow clarifier or a sludge blanket clarifier.
The latter term can be misleading, however, since not all such units operate with a
sludge blanket. A diagram of a unit that does not operate with a sludge blanket is
shown in Figure 8-4. Gulp and Gulp (1) recommend the system of separate tanks
because that system allows separate control of each part of the process. They point
out that such a system has greater flexibility in points of addition for chemicals. They
also point out that in the integrated units with a sludge blanket there may be
difficulty in controlling the blanket height and the blanket may become anaerobic,
leading to poor phosphorus and solids removal. Pilot studies at other locations have
shown that excellent solids removal can be obtained with sludge blanket equipment,
but that blanket instability could be a problem. On the other hand, pilot upflow units
designed to operate without a sludge blanket have proved to be very effective at
Washington, D. C. (2). Results available at this time indicate that both the system of
separate tanks and the integrated units without sludge blankets should be considered
for design of new plants.
The Tahoe plant is an excellent example of a system with the first stage consisting of
separate tanks for each unit process. At the present time the plant must be relied
upon heavily as a guide to design of such systems. Description of the equipment and
some design parameters are given by Gulp and Gulp. The rapid mix tank which is
located in one corner of the flocculator has about a 30 second residence time at
design flow. (The plant usually runs with a daily average flow of about one third of
design flow.) Mixing is accomplished with a vertical shaft mixer. The flocculator is a
8 - 10
-------
WASTEWATER
LIME
D
cu
Q
CO
Q
LLJ
_J
O
O
LLJ
EFFLUENT
u
SLUDGE
FIGURE 8-4 TYPICAL UPFLOW CLARIFIER
8 - 11
-------
square tank with a depth of 8 ft. At design flow the residence time is 4.5 minutes.
The flocculator is provided with air agitation, but operating experience has shown that
this agitation is not necessary. The settler is circular with a center inlet. The depth is
^
10 ft and the design overflow rate is 950 gpd/ft . There are two sludge pumps, one a
variable-speed centrifugal, and the other a positive displacement Moyno. Provision has
been made to return part of the sludge to the rapid mix tank. The concept of
returning sludge to the section of the clarification equipment where precipitation is
taking place is called solids contact. The objective is to hasten precipitation and to
obtain larger precipitate particles which will settle well.
The ease with which flocculation occurs at Tahoe would suggest that a flocculator for
this service presents no design problems. Experience is, however, limited. Pilot testing
would be very desirable to determine the flocculating characteristics of a particular
wastewater. Jar tests may be of some value. Unfortunately, the effect of solids
contact, which experience at Tahoe and elsewhere indicates is beneficial, is difficult to
duplicate in a jar test. If jar tests indicated good flocculating characteristics without
benefit of solids contact, however, flocculation should be as good or better with solids
contact. When pilot studies are carried out, provision should be made for solids
contact.
Comments concerning the effect of solids contact in jar tests also apply to settling
rates determined from these tests. Rapid settling without benefit of solids contact
would be a strong indication that settling with solids contact would also be rapid.
Clarification equipment of the integrated type for use in municipal water treatment is
manufactured by a large number of companies. Units of this type used for treatment
of wastewater have been essentially of the same design as used for municipal water
supplies. The basins are usually circular with the rapid mix and flocculating sections
located at the center of the settler. Agitation for mixing and flocculating as well as
power for sludge scrapers is provided at the center of the settler. Solids contact is
usually provided, and may be carried out with best control by external circulation of
sludge to the mixing section. In some designs, however, internal recirculation is used.
There is considerable variation in the geometry and complexity of the mixing and
flocculating sections depending upon the manufacturer. The unit shown in Figure 8-4
is a relatively simply type. Units from three different manufacturers were used in pilot
work at Lebanon, Ohio and Washington, D. C. These and others at additional locations
have performed effectively. However, in some cases with operation of a sludge blanket,
difficulties have been encountered. A sludge blanket can be eliminated by avoiding
designs in which the wall separating the center compartments from the concentric
settler reaches close to the bottom of the settler. The design in Figure 8-4 prevents
sludge blanket formation.
There has been little effort to modify this type of equipment from use in water
treatment to use with wastewater. Experience at Tahoe would suggest, for example,
that less flocculation time, and probably less agitation, is required than is generally
provided. A unit such as that shown by Figure 8-4 may have about 40 minutes
flocculation time at design capacity for water treatment, nearly nine times that at
8 - 12
-------
Tahoe. Whether modifications in such equipment would be worthwhile remains to be
seen.
If a designer chooses to use equipment that is on the market, he does not have
complete flexibility in specifying the design. Fortunately, equipment satisfactory for
water supply treatment has proved satisfactory for wastewater, except that a lower
settler overflow rate is necessary. The single-stage treatment system at Lebanon, Ohio
^y
was operated at a constant overflow rate of 1,440 gpd/ft . This proved satisfactory
even with a sludge blanket. At Washington, D. C., tests were made with settler rates as
high as 1,950 gpd/ft2 although the average rate was 1,120 gpd/ft2. For design, a peak,
^
dry weather overflow rate of 1,200 to 1,400 gpd/ft , with the average somewhat lower,
appears reasonable.
8.4.2 Recarbonation
In a two-stage system recarbonation follows the first stage settler. In a single-stage
system and following the second stage of a two-stage system, pH reduction by
recarbonation may be necessary to make the effluent suitable for filtering or for
discharge. An excellent discussion of all aspects of wastewater recarbonation is given
by Gulp and Gulp (1). The reader is referred to that publication, as well as one by
Haney and Hamann (5), especially for information on sources of carbon dioxide.
Sources include stack gas from sludge incinerators and lime recalciners, CC^ generators,
and commercial liquid
The equipment used for contacting the CC^ and water may be simply a tank with the
gas being bubbled through the water. This is the system used at the Tahoe plant
where the height of water over the CC>2 source is 8 ft. Pilot studies at Washington,
D. C., with a tank depth of 1 1 ft and a turbine mixer to reduce bubble size and
distribute bubbles, showed almost 100% absorption of CCK-
When considering residence time, the recarbonation tank following the first stage of a
two-stage system must be differentiated from the recarbonation tank used just to
reduce effluent pH. The residence time of the recarbonator between the first and
second stages is not particularly important. At Tahoe the recarbonation tank is only
large enough to give a 5 minute residence time at design flow. In the Washington,
D.C. studies, the residence time was 15 minutes. A problem that may arise when
recarbonating wastewater is foam formation. This is most pronounced when the flow
of bubbles is concentrated in parts of the recarbonation tank. The CC<2 distributing
system should cover as well as possible the whole lateral area of the tank bottom. This
would be especially true if the residence time were as low as 5 minutes.
In the recarbonation tank used just for effluent pH adjustment, sufficient residence
time must be allowed for completion of reactions taking place. Gulp and Gulp
recommend 15 minutes. In the Tahoe plant only 4 minutes is provided at design flow,
but the water flows to storage ponds with a much longer residence time before the
water is filtered. Just as in the case of recarbonation between stages of a two-stage
system, good bubble distribution is important for prevention of excessive foaming.
- 13
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Since one reason for using a two-stage system is to prevent a high calcium
concentration in the effluent, the pH to which the water should be recarbonated
between stages is that for maximum conversion of the calcium to CaCOo. Where lime
recovery is practiced, maximum formation of CaCOo is also desirable because it results
in maximum CaO production. The optimum pH is about 10. At Tahoe, for example,
10.3 was selected because tests showed that pH to give maximum CaCCK production.
The pH to be selected when the objective of recarbonation is just pH reduction, will
depend on a number of factors. An effluent standard may determine this pH. If
stabilizing the water to prevent post precipitation is the objective, the pH may be
determined by consideration of the Langelier Index. If, as in the case of Tahoe, a
treatment step such as activated carbon adsorption is to follow lime treatment, this
will determine pH. In the latter case a pH of about 7.5 has usually been selected.
The CO 2 dose requirements can be calculated from the chemical reactions taking place
and a knowledge of the concentrations of the various forms of alkalinity in the water.
For design of a new plant, alkalinity data must be obtained from samples of liquor
taken from jar tests run at the same pH values as planned for the plant. Results must
be considered very approximate, but they do help in sizing CC»2 feeding equipment.
Gulp and Gulp discuss the calculations in detail. It was found in pilot plant work at
Washington, D. C. that the CO^ dose for recarbonation between the stages of a
two-stage treatment system can be calculated reasonably well by considering the
reduction in calcium concentration that occurs. This reduction is due to CaCOo
7
formation with the CO^ coming from the CC^- The calcium content before
recarbonation can be determined satisfactorily from jar tests run at the desired
operating pH. The calcium content after recarbonation can also be determined
approximately by jar test. Results from several pilot plants indicate the soluble calcium
content at the pH of minimum solubility should be about 40 mg/1 as Ca. This figure
is probably just as good for calculations as a figure obtained from a jar test. The CC>2
dose in mg/1 is then equal to:
/44\
(Ca reduction in mg/1) — .
\40/
A safety factor of about 20% should be added to the calculated dose to compensate
for inefficiency in absorption.
8.4.3 Second Stage Flocculation and Sedimentation
In two-stage treatment, recarbonation is followed by the second stage flocculation and
settling. Experience with equipment for the second stage of treatment with wastewater
is limited. Two very different systems have been tested. The simplest is the system at
Tahoe. In the Tahoe plant the equipment used is just a longitudinal settler with a 30
O
minute detention time and a 2,400 gpd/ft overflow rate at design flow. No
flocculation equipment is provided. Gulp and Gulp (1) refer to the settler as a reaction
and settling basin. It has been found at Tahoe that a significant part of the
recarbonation reaction actually occurs in the settler. This is shown by the pH decrease
8 - 14
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of about 0.7 unit that occurs between the recarbonation tank and the settler outlet.
The clarification that results at Tahoe is unusually good for the simplicity of the
equipment.
Jar test and pilot plant work at Washington, D. C. indicated that flocculation after
addition of a flocculating aid was required to obtain satisfactory clarification of that
water. The pilot tests were run in upflow equipment of the integrated type discussed
earlier. Without a flocculating aid, fine CaCOo precipitate escaped the settler, even
^) "\ I
at a settler overflow rate as low as 1,100 gpd/ft . Using 5 mg/1 of Fe , good
^
operation was observed at overflow rates as high as 1,950 gpd/ft . Hydraulic
limitations prevented testing at higher rates. Apparently, the good contact between
solids and water during flocculation was beneficial in reducing CaCO^ supersaturation
in the effluent. Precipitation in the filters following the second stage did not occur,
and no pH reduction was required.
In this work the sludge from the second stage was returned to the mixing section of
the first stage to serve as a weighting agent. In waters of low alkalinity or expected
high loads of biological solids, this procedure should be considered.
It is difficult to recommend the amount of flocculation to provide and the overflow
rate to use for second-stage treatment. Jar tests would be of little value because of the
complication of first raising the pH with lime and then lowering with CO^- Pilot
testing may be desirable. Such a pilot system would have to include the first stage
equipment and recarbonation equipment, although these could be of crude design. In
the absence of pilot tests to indicate otherwise, flocculation should be provided in the
second-stage system. If filtration is included, a peak dry weather overflow rate of up
to 2,000 gpd/ft^ in the settler may be used. If filtration is not provided, a somewhat
lower overflow rate should be chosen to assure reasonable effluent quality during periods
when an upstream part of the system is not operating properly.
8.4.4 Filtration
The last step in a complete lime treatment system is filtration. Although there are a
variety of filter types that could be used, essentially all the recent test work has been
done with downflow filters of multimedia type. These are essentially the same in
design as a rapid sand filter, except that the media are graded from coarse at the filter
surface to fine at the filter outlet. This is accomplished by using media of different
densities with the largest particles being composed of the least dense material. The
result is a filter bed which has far more capacity for removing suspended solids than
an ordinary sand filter, without impairing effluent clarity. Both gravity and pressure
filters have been used. Each has certain advantages. Generally, however, pressure filters
are more appropriate at smaller plants.
Two media combinations have been tested extensively. These are dual-media of coal
and sand, and tri-media of coal, sand, and garnet. Filters used at Lebanon, Ohio
following a single-stage system (3) and filters at Washington, D.C. following a
two-stage system (2) were of the dual-media type. Both contained 18 in. of anthracite
8 - 15
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coal over 6 in. of sand. Average particle sizes of the media at Lebanon were 0.75 mm
^
and 0.46 mm; at Washington, D.C., 0.9 mm and 0.45 mm. At 2 gpm/ft the average
run length at Lebanon was about 90 hours using 8 ft of water as the terminating
pressure loss. At Washington, D. C., filter run length averaged about 50 hours for an
average rate of 3.4 gpm/ft ^ and a terminating pressure loss of 7 ft of water. High
clarity waters were obtained in each case. Backwash in both cases was carried out at
•"}
20 gpm/ft . Backwash time was 5 minutes at Lebanon and 10 minutes at Washington,
D. C. In the latter case, a surface wash was included.
These run length figures give a rough idea of the results that may be expected at
other locations. For maximum run length at a given product quality, the depths of
media and their particle sizes should be optimized. Although this was not done for the
above-mentioned filters, pressure loss distribution in the filters at Lebanon as reported
by Berg, et al. (3) indicates that good solids storage in the anthracite was being
obtained. Small pilot filters could be used for optimization of the media.
A reasonable design rate for gravity-flow dual-media filters appears to be about 3
gpm/ft . Higher rates could result in inconveniently short runs during periods of high
solids load. Rates much lower than 3 gpm/ft ^ result in excessive filter costs, not
justified by the longer filter runs. Very long filter runs could result in backwash
problems from biological activity in the filters. Recommendation of a 3 gpm/ft^ rate is
based upon having good operation of the second-stage settler. Frequent settler upsets
must be avoided.
Tri-media filters are used at the Tahoe plant. These are pressure filters which at design
flow operate at 5 gpm/ft . Each bed holds 3 ft of mixed coal, sand, and garnet as
supplied by Neptune Microfloc. Two filters are used in series, and are backwashed in
series, usually when the head loss reaches 16 ft of water. Run length has varied from
4 hours during heavy solids load to about 60 hours under good conditions. Backwash
is carried out at 15 gpm/ft^. The backwash water is reprocessed through the lime
treatment system. More details are given by Gulp and Gulp (1).
From the available data, it is difficult to give precise criteria for choosing between
dual-media or tri-media filters following lime treatment. The use of garnet allows for a
smaller particle size at the bottom of the filter than is possible with sand. In case of
floe weakness, the tri-media filter offers, therefore, more protection from turbidity
breakthrough. The cost, however, is slightly higher. Where the tri-media fill is not used,
filters can still be backwashed at the onset of turbidity breakthrough. Experience with
dual-media filters on lime treated water has not shown sudden breakthrough to be an
important problem. Where water of the highest clarity is required, tri-media filters may
be of most value.
8.4.5 Sludge Handling
The sludge from a lime treatment system may be handled in two general ways. It may
be thickened, dewatered, and disposed of or it may be thickened, dewatered, and
8 - 16
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recalcined to recover lime for reuse. For small plants, recalcination will not be economically
competitive with disposal and should not be considered unless there are restrictions
on disposal which make that alternative difficult. It can be assumed generally that
for plants over 10 mgd, recalcination of lime sludge will be practical. Recalcination
may also be practical at plants somewhat smaller than 10 mgd, depending on the cost
of purchased lime and other local conditions.
Data given earlier indicate that the volume of sludge from lime treatment will vary
from 1.5% to several percent of feed volume. Sludge concentration will probably be in
the range of 1% to 5%. The actual weight of sludge will vary with the chemical
composition and amount of suspended solids in the feed water. Values have been
observed in the range of 4 to 7 lb/1,000 gal. Sludge production from the first stage of
a two-stage system or from a single-stage system can be estimated approximately by
weighing the dried sludge from jar tests run at the planned operating pH. Additional
sludge produced in the second stage can be assumed to be the CaCC^ formed from
calcium concentration reduction during recarbonation. Calcium content before
recarbonation can be obtained from the above-mentioned jar tests and after recarbonation
can be assumed to be 40 mg/1 as Ca.
Design information for the Tahoe sludge handling system is reported by Gulp and
Gulp. (1) The reader is referred to that publication for further information. The Tahoe
plant is the only one which includes recalcination of sludge from the lime treatment
of wastewater, and that has been operated for a long period of time. Paramenters from
that system can be used as a rough guide for design. The gravity thickener has a
bottom scraper mechanism and is 8 ft deep. Overflow from the thickener is returned
to either the primary clarifier or the first stage of lime treatment. The design solids
loading is 200 Ib/day/ft^ and the design overflow rate is 1,000 gpd/ft^. Some
preliminary data from pilot work at Washington, D. C. suggest these loadings are high.
The Tahoe equipment was sized, however, to handle a volume of sludge equal to
about 9% of the plant design flow. The actual volume of sludge should usually be less
than one third of that volume. At Washington, D. C., recalcined lime was found to
produce a sludge which thickened significantly faster than sludge from virgin lime.
If lime is not to be recovered from the sludge, the underflow from the thickener can
be placed directly on drying beds for final dewatering, or it can be dewatered by
centrifuge or vacuum filter.
If sludge recalcination is planned, the thickened sludge would be dewatered, by
centrifuging or filtering and fed to the calcining furnace. The centrifuge may have
advantages over filters for this purpose, such as the ability to separate phosphate
sludge from CaCO^. Lime from the furnace is stored for reuse. It, along with any new
lime, must be slaked before use. Sludge can be pumped from the thickener to the
centrifuge. Cake from the centrifuge must be transported by conveyor. At Tahoe the
centrate is sent to the primary clarifiers. It has been found at Tahoe that, by
operating the centrifuge with less than maximum solids capture, much of the
precipitated phosphorus can be retained in the centrate. This results in a higher
- 17
-------
quality lime. This mode of operation requires a second centrifuge to remove the
phosphorus rich solids from the centrate. These solids may have value for use in
fertilizer.
The Tahoe dewatering and recalcining system is described by Gulp and Gulp. The
centrifuge is a 24 by 60 in. concurrent flow type. The calciner is a 14 ft-3 in. diameter,
6 hearth furnace operated at a top temperature of 1,850° F. Other types of
furnaces have been used in water treatment plants and these should also be applicable
to the sludge from wastewater treatment. The reader is referred to equipment
manufacturers for further information about centrifuges and furnaces.
8.4.6 Control of Lime and Carbon Dioxide Feed
The feeding of lime to a lime softening system can be controlled in a number of
ways. For the wastewater application, the most appropriate appears to be control of
pH, with the pH value being selected for good suspended solids removal or to meet a
phosphorus removal requirement. Where flow equalization is employed, pH sensing
alone should be sufficient. Where there is substantial diurnal variation in flow, better
control can be maintained by flow proportional-pH control. Systems are available
for this type of control.
Carbon dioxide dose is also controlled by pH. As in the case of lime feed, flow
proportional-pH control would be preferred where there is diurnal variation in flow.
There is a tendency for pH electrodes to become coated with precipitates and lose
sensitivity. The precipitate can be dissolved with acid. In some instances, however, it
has been found satisfactory to use manual control because of difficulties with the pH
control system.
8.4.7 Scale Formation on Equipment
Because the water in a lime treatment system is supersaturated to some degree with
CaCOo and other precipitating substances, there is a tendency for scale to form on
equipment and pipe surfaces. The problem is particularly serious with the lime slurry
from the slaker.. This can quickly plug the slurry line to the rapid mix tank. There
should be easy access to all parts of this line for cleaning. At one small plant a
flexible hose was used to feed lime slurry. By periodically flattening or flexing the
hose, the scale was removed. The mechanical mixer in the rapid mix tank will also
become scaled and must be cleaned.
Scale can also form in sludge lines and the effluent line from the first-stage clarifier.
It is recommended that open troughs be used wherever possible. Provision should be
made for cleaning lines when open conduits are not possible.
Possible scaling of filters has already been mentioned. Recarbonation before the filters
will minimize scale formation.
8 - 18
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8.5 Capital and Operating Costs
Because of the short history of lime treatment of wastewater, there is a scarcity of
capital and operating cost information. Costs are available for the Tahoe plant and are
reported by Gulp and Gulp (1). Capital cost for the 7.5 mgd lime treating facility was
$1,115,000 and cost of the filters was an additional $705,000. For 1969 the estimated
operating costs exclusive of equipment amortization were 7.3^/1,000 gal. for lime
treatment without filtration and 2.8^/1,000 gal. for filtration. Amortization of
equipment, based on costs adjusted to the 1969 national average, interest of 5% for 25
years, and the assumption of the plant operating at full capacity, would add
2.7^/1,000 gal. for treatment without filtration and 1.8 #1,000 gal. for filtration. For
the plant operating at full capacity the total cost of operation would then be
10.0#1,000 gal. without filtration and 14.6#1,000 gal. including filtration. The
fraction of the cost resulting from amortization is a substantial 31% even with the
assumption of full capacity. There is an obvious need to keep equipment size to a
minimum. Flow equalization deserves strong consideration when planning for lime
treatment of secondary effluent to minimize the initial cost of new equipment.
Additional costs for lime treatment based on information from Tahoe and other
sources have been reported by Smith and McMichael (6). Tables 8-2 and 8-3 show
Table 8-2
CAPITAL COST OF LIME TREATING FACILITIES
Cost ($)
Treatment Plant Size (mgd)
U) K) 100
Single-Stage without Filtration 100,000 1,200,000 5,500,000
Two-Stage without Filtration 160,000 1,500,000 7,900,000
Dual-Media Filtration 110,000 510,000 2,300,000
Table 8-3
TOTAL COST FOR LIME TREATMENT OF WASTEWATER
Cost (#1,000 gal.)
Treatment Plant Size (mgd)
LP 1P_ 100
Single-Stage without Filtration 13 7 4
Two-Stage without Filtration 16 9 6
Dual-Media Filtration 8 3 1.4
costs based largely on their results. Capital costs have been updated to December,
1970. Amortization is at 6% for 25 years. For 1 mgd plants, recalcination equipment is
- 19
-------
not included. All lime would be purchased. The applicability of recalcination was
discussed earlier. Component costs making up the total cost figures represent an
average for the whole country. Local conditions may cause significant deviation from
these values.
8 - 20
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8.6 References
1. Gulp, R. L., and Gulp, G. L., Advanced Wastewater Treatment, Van Nostrand
Reinhold Company, New York (1971).
2. O'Farrell, T. P., and Bishop, D. F., "Lime Precipitation in Raw, Primary, and
Secondary Wastewater", Presented at the 68th National Meeting of AIChE,
Houston, Texas (March 1971).
3. Berg, E. L., Brunner, C. A., and Williams, R. T., "Single-Stage Lime Clarification of
Secondary Effluent", Wat. and Wastes Eng., 7:3, p 42 (1970).
4. Buzzell, J. C., and Sawyer, C. N., "Removal of Algal Nutrients from Raw
Wastewater with Lime", JWPCF, 39:10, Part 2, p R16 (1967).
5. Haney, P. D., and Hamann, C. L., "Recarbonation and Liquid Carbon Dioxide",
JAWWA, 61:10, p 512 (1969).
6. Smith, R., and McMichael, W. F., "Cost and Performance Estimates for Tertiary
Wastewater Treating Processes", Robert A. Taft Water Research Center, Report No.
TWRC-9 (June, 1969).
8 -21
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Chapter 9
PHOSPHORUS REMOVAL BY MINERAL
ADDITION TO SECONDARY EFFLUENT
9.1 Description of Process
Alum and to a lesser extent iron salts have been evaluated as precipitants for phosphorus
in effluents from conventional treatment plants. Both batch studies and continuous-flow
systems up to 2.5 mgd have successfully reduced phosphorus to very low levels.
Equipment typically consists of mixers, flocculators, settlers, and filters. Some work
has been done, however, with the chemicals added directly to filters. Specific studies
are discussed later and provide additional information on equipment commonly used.
9.2 Summary of Design Information
An alum dosage of about 200 mg/1 is required for phosphate removal from typical
municipal wastewater while dosages of 50 to 100 mg/1 are sufficient for effluent
clarification. Iron salts, while successful in precipitating phosphate, have found little
application because of residual iron remaining in the treated water. An A1:P molar
ratio of 1:1 to 2:1 is required. The optimum pH for alum treatment is near 6.0
while for iron the optimum pH is near 5.0 (1). The pH of high alkalinity
waters may be reduced either by using high dosages of alum or by adding
supplementary dosages of sulfuric acid. Anionic polyelectrolytes are useful with
alum in improving settleability of floe. Settling alone will reduce residual P to
about 1 mg/1.
If higher removals are required, filtration must be employed. With proper operation,
residual P may be reduced to less than 0.1 mg/1. Multimedia filters are preferred over
sand filters because of the extended length of run. Microstrainers have not been
successful in removing alum floe.
/^
Surface overflow rates for clarifiers have ranged from 580 to 1,440 gpd/ft , but for
consistent removal of phosphate, the lower end of the range is preferred. Filtration
rates of 2 to 5 gpm/ft are used.
No process has yet been developed for successfully recovering alum from sludge for
reuse although both alkaline and acid regeneration were investigated (2, 3). The sludge
can be dewatered by conventional methods. Because dewatering may be difficult,
design of such facilities should be conservative.
Costs for tertiary treatment for phosphate removal range from 6 to 15^/1,000 gal.
depending on plant size and degree of treatment required.
9 - 1
-------
9.3 Laboratory and Pilot Studies
9.3.1 South Lake Tahoe (3, 4, 5, 6)
Liquid alum was used to precipitate phosphate in a 2.5 mgd tertiary plant at South
Lake Tahoe, California. The treated wastewater was applied to multimedia filters for
solids removal without prior sedimentation.
Chemical feed was proportioned automatically to flow, and flash mixing was provided
by an in-line mixer in the filter influent line. The filter media consisted of layers of
anthracite coal, graphite, fine garnet and coarse garnet overlaying a supporting bed of
gravel. Two filters were operated in series at 5 gpm/ft . Backwash with secondary
plant effluent was initiated either by headless or by a rise in turbidity of the effluent.
Polyelectrolyte was used as a filter aid to control the depth of penetration of floe into
the filter beds. With an alum dosage of 200 mg/1, 8 to 9 mg/1 P in the secondary
plant effluent was reduced to 0.03 to 0.3 mg/1 P in the filter effluent. Turbidities
were correspondingly reduced from 30 to 70 JTU to 0.2 to 3.0 JTU.
Efforts to regenerate alum from the sludge by both acid and alkaline treatment proved
to be uneconomical and a lime treatment process was developed as a substitute for the
alum process. Lime could be economically recovered for reuse and had the added
advantage of producing a high pH which facilitated removal of nitrogen by ammonia
stripping.
Total costs for phosphate removal only, as estimated in 1966, are shown in Table 9-1.
Table 9-1
1966 ESTIMATED COSTS FOR TERTIARY
PHOSPHATE REMOVAL WITH ALUM
Plant Capacity (mgd) Total Cost (jj 1,000 gal.)
2.5 11.7
10 8.5
50 7.1
100 6.6
200 6.3
9.3.2 Nassau County (7, 8)
Nassau County, New York, is evaluating treatment methods for its wastewater to
provide a product water which is suitable for ground injection as a barrier to salt
water intrusion. A pilot plant has been operated at 400 gpm producing a water which
in most respects exceeds USPHS drinking water standards. Effluent from high rate
activated sludge is pumped into a 40 ft diameter by 14 ft deep clarifier where alum
and poly electrolyte are added. Sludge is recirculated into the coagulation zone to
improve efficiency of floe formation. The treated wastewater then passes downward
through a flocculation zone and upward through a clarification zone to orifice type
9 - 2
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collectors. The clarified wastewater passes to 2 dual-media filters in parallel operated at
3.0 gpm/ft2. Each filter consists of a 36 in. bed of No. 1-1/2 anthracite (effective size
of 0.90 mm) above a 12 in. layer of sand (effective size of 0.40 mm). Filter backwash
includes air scour, surface wash and high and low-rate backwashing. The filter effluent
passes through 4 granular carbon contactors operating in series to remove organics.
Typical operating results with about 200 mg/1 alum are shown in Table 9-2.
Table 9-2
NASSAU COUNTY PERFORMANCE DATA
Turbidity, JTU
Color, Pt-Co Units
COD, mg/1
P, mg/1
Influent
20 - 25
20 - 40
78
Effluent
0.2 - 1.5
none
13
0.1 - 1.0
Costs have been estimated, for separate tertiary treatment only, at plant sizes of ], 10,
and 100 mgd as shown in Table 9-3.
Table 9-3
NASSAU COUNTY COST ESTIMATES
Process
Coagulation
Filtration
Carbon adsorption
Operating labor
Total Cost
9.3.3 Chicago, Hanover Park (9)
(1)
4.9
1.8
6.3
13.0
28.0
41.0
Plant Capacity (mgd)
(10) (100)
(Costs, H 1,000 gal)
3.5
1.1
4.5
9.1
5.6
14.7
3.2
1.0
4.0
8.2
1.8
10.0
Alum was used to coagulate solids and precipitate phosphate from secondary effluent
at Chicago's Hanover Park tertiary treatment plant. Treatment consisted of
coagulation-sedimentation in circular clarifiers followed by either rapid sand filtration
or microstraining for polishing. Average annual flow was 1.5 mgd.
Combined coagulation plus sand filtration removed 34, 58, and 82% P at dosages of
38, 60, and 140 mg/1 of alum, respectively. Combined coagulation plus microstraining
effected a lower phosphate removal of 23% at an alum dosage of 38 mg/1. In general,
9 - 3
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alum coagulation under the test conditions gave little or no improvement in suspended
solids removal over that obtained by sand filtration of uncoagulated secondary effluent.
The microstrainer did not significantly remove alum-coagulated solids.
9.3.4 Dallas (10)
Dallas, Texas, is operating a demonstration project designed to provide necessary design
data for upgrading of the Dallas-White Rock trickling filter plant. Phosphorus removal
facilities include a solids contact upflow clarifier unit and two multimedia filters. The
clarifier has a rise rate of 0.5 gpm/ft2 at the design flow of 100 gpm and a detention
time of 4 hours. Both filters contain^ anthracite over sand media. One is a standard
gravity filter. The other is a biflow type, with feed water entering at both the top and
bottom of the bed. Filtered water is removed by a manifold-lateral collector located at
the mid-depth of the bed. The gravity filter is designed for 50 gpm at a filtration rate of
0 *")
4 gpm/ft . The biflow filter is designed for 100 gpm at a filtration rate of 8 gpm/ft , or
^
4 gpm/ft^ for both the bottom and top of the bed.
The pilot plant process has not yet been optimized for phosphorus removal, however,
limited data from two short test periods on trickling filter and activated sludge effluents
are shown in Table 9-4.
Table 9-4
REMOVAL OF TOTAL ORTHOPHOSPHATE BY COAGULATION-CLARIFICATION
IN THE DALLAS DEMONSTRATION PLANT
Trickling Filter Activated Sludge
Effluent Effluent
Lime Dose, mg/1 CaO 128 133
FeCl3 Dose, mg/1 Fe3+ 25 56
Soluble Phosphate In, mg/1 P 6.5 7.0
Soluble Phosphate Out, mg/1 P 0.43 0.21
% Soluble P Removed 93.5 97.0
Test Period, days 12 2
9.3.5 Dayton (11)
Dayton, Ohio, used alum in a 310 gpm pilot plant to remove phosphate and suspended
solids from trickling filter effluent. The rectangular sedimentation basin was designed for
105 minutes detention and an overflow rate of 1,014 gpd/ft . Two filters designed for
^
loadings up to 8 gpm/ftz were used for polishing effluent. One contained 30 in. of sand
while the other used 20 in. of anthracite overlaying 10 in. of sand. The sand had an
effective size of 0.40 to 0.45 mm with a uniformity coefficient of 1.35 to 1.70. The
anthracite had an effective size of 0.85 to 1.20 mm with a uniformity coefficient of 1.35
to 1.80. Both air and water were used in backwashing.
Alum with activated silica or anionic polyelectrolyte as coagulant aids achieved
•3
phosphate removals at a dosage rate of 5 Ib alum/lb PO^ down to residuals of 4
9-4
-------
mg/1 PC>4 in filter effluents (89% removal of phosphate). Further removals down to
1 to 2 mg/1 PC>4 required in excess of 200 mg/1 alum. The clarifier was operated at
overflow rates of 655 to 1,035 gpd/ft2.
The filters removed 30 to 70% of the applied phosphate load. The dual media filter
produced an effluent comparable to that from the sand filter with filter runs 2 to 3
times the length of the sand filter runs.
Costs to provide 90% phosphate removal with alum for the present 50 mgd flow in a
75 mgd plant were estimated at 6.5 to 8.2^/1,000 gal.
9.3.6 Oxidation Pond Effluents
At Lancaster, California, a research project was initiated to renovate wastewater for use
in recreational lakes (12). Existing treatment consisted of primary sedimentation
followed by 45 to 60 days detention in oxidation ponds. For coagulation, laboratory
jar tests showed alum to be more effective on a cost basis than either lime or ferric
sulfate. Phosphate residuals were reduced to 1.1, 0.2, and 0.01 mg/1 P with alum
dosages of 150, 200, and 300 mg/1, respectively, at an optimum pH near 6.0. Pilot
studies showed that both sedimentation and air flotation would provide satisfactory
separation of most of the precipitated phosphates, however, filtration was necessary to
obtain the required low residuals of less than 0.15 mg/1 P.
Final specifications for a full scale process included coagulation with about 300 mg/1
alum and 20 minutes of flocculation with a conventional paddle flocculator,
sedimentation in a horizontal clarifier with an overflow rate of 400 gpd/ft , and
filtration through a dual media gravity filter consisting of 18 in. of 0.5 mm anthracite
over 8 in. of No. 20 sand.
Capital costs of a 0.5 mgd facility were estimated' at $150,000 with operating costs of
$184/10"gal. The total cost of capital, operation, and maintenance for a 3 mgd
facility was estimated at S150/106 gal.
Alum was recently evaluated by Shindala and Stewart (13) for phosphorus removal and
polishing of waste stabilization pond effluents in Mississippi. Jar tests were used in the
evaluation. Using a minimum of 90% removal of phosphorus and 70% removal of COD
as criteria, the optimum alum dosage was 85 mg/1 at pH 5.5. With higher doses, up to
97% phosphorus removal and 85% COD removal were observed. Ferric chloride and
ferric sulfate were also tried but left a residual iron color in the supernatant.
9.3.7 Moving Bed Filter (14)
Alum and polyelectrolyte were used to treat trickling filter effluent in a proprietary
moving bed filter at Liberty Corner, New Jersey. The device, a product of
Johns-Manville Products Corp., used a bed of sand which is periodically moved upward
in an inclined filter vessel. Clean sand is fed in at the bottom of the filter and
floe-laden sand is removed mechanically at the top as required by pressure drop,
9 - 5
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leaving a fresh filtration surface. The wastewater being filtered passes down through
the bed and flows through screens at the sides of the filter body.
A major factor controlling phosphate removal is the molar ratio of A1:P. At an alum
dosage of 200 mg/1 and influent total phosphorus concentrations on the order of 25
to 28 mg/1 as P (A1:P ratio = 0.6 to 0.7), the total phosphorus removal efficiency
averaged 90%. With lower total phosphorus concentrations (A1:P ratios = 1.2 to 2.6)
removal efficiencies averaged 95% and reached as high as 99%. The alum treatment
also removed about 70% of suspended solids leaving an average of 15 mg/1 in the
effluent.
Capital costs for a 1 mgd moving bed filter are estimated at $264,000. Total operating
costs for such an installation using 200 mg/1 alum are estimated at 12^/1,000 gal.
treated.
9 - 6
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9.4 References
1. Recht, H. L., and Ghassemi, M., "Kinetics and Mechanism of Precipitation and
Nature of the Precipitate Obtained in Phosphate Removal from Wastewater Using
Aluminum (III) and Iron (III) Salts", Water Pollution Control Research Series
17010EKI, Contract 14-12-158, USDI, FWQA (April, 1970).
2. Farrell, J. B., Salotto, B. V., Dean, R. B., and Tolliver, W. E., "Removal of
Phosphate from Wastewater by Aluminum Salts with Subsequent Aluminum
Recovery", Chem. Eng. Prog. Symp. Series, 64:90, p 232 (1968).
3. Slechta, A. F., and Culp, G. L., "Water Reclamation Studies at the South Tahoe
Public Utility District," JWPCF, 39:5, p 787 (1967).
4. Culp, R. L., "Wastewater Reclamation by Tertiary Treatment", JWPCF, 35:6, p 799
(1963).
5. Culp, R. L., and Roderick, R. E., "The Lake Tahoe Water Reclamation Plant",
JWPCF, 38:2, p 147 (1966).
6. Culp, R. L., "Wastewater Reclamation at South Tahoe Public Utilities District",
JAWWA, 60:1, p 84(1968).
7. Rose, J. L., "Injection of Treated Waste Water into Aquifers", Wat. & Wastes Eng.,
5:10, p 40(1968).
8. Rose, J. L., "Advanced Waste Treatment in Nassau County, N. Y.", Wat. & Wastes
Eng., 7:2, p 38 (1970).
9. Lynam, B., Ettelt, G., and McAloon, T., "Tertiary Treatment at Metro Chicago by
Means of Rapid Sand Filtration and Microstrainers", JWPCF, 41:2 (Part 1), p 247
(1969).
10. Graeser, H. J., and Haney, P. D., "Dallas Builds Center to Study Wastewater
Reclamation", Wat. & Wastes Eng., 5:12, p 34 (1968).
11. Tossey, D. F., Fleming, P. J., and Scott, R. F., "Tertiary Treatment by Flocculation
and Filtration", Jour. SED, ASCE, 96:SAI, p 75 (1970).
12. Dryden, F. D., and Stern, G., "Renovated Waste Water Creates Recreational Lake",
Environmental Science & Tech., 2:4, p 268 (1968).
13. Shindala, A., and Stewart, J. W., "Chemical Coagulation of Effluents from Municipal
Waste Stabilization Ponds", Wat. & Sew. Works, 118:4, p 100 (1971).
14. "Phosphorus Removal Using Chemical Coagulation and a Continuous Countercurrent
Filtration Process", Water Pollution Control Research Series 17010EDO USDI,
FWQA (June 1970).
9 -7
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CHAPTER 10
STORAGE AND FEEDING OF CHEMICALS
10.1 Aluminum Compounds
The principal aluminum compounds that are commercially available and suitable for
phosphorus precipitation are alum and sodium aluminate (1). Both are available in either
liquid or dry forms. Alum is acidic in nature while sodium aluminate is alkaline and
this may be an important factor in choosing between them. Aluminum chloride should
be considered, and is discussed herein.
10.1.1 Dry Alum
10.1.1.1 Properties and Availability
The commercial dry alum most used in wastewater treatment is known as "filter alum"
and has the approximate chemical formula A^CSO^ 14^0 and a molecular weight
of 594. Alum is white to cream in color. The pH varies between 3.0 and 3.5 in aqueous
alum solutions having concentrations between 1% and 10%. Commercially available grades
and their corresponding bulk densities and angles of repose are given in Table 10-1:
TABLE 10-1
AVAILABLE GRADES OF DRY ALUM
Grade Bulk Density Angle of Repose
Ib/cu ft
Lump 62 to 68 varies
Ground 60 to 71 43°
Rice 57 to 61 38°
Powdered 38 to 45 65°
Each grade has a minimum aluminum content of 17%, expressed as A^Oj. This
corresponds to a 9% concentration as aluminum. Viscosity and solution crystallization
10-1
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temperatures are included in the subsequent section on liquid alum.
Since dry alum is only partially hydrated, it is slightly hygroscopic. However, it is relatively
stable when stored under the extremes of temperature and humidity encountered in the
contiguous United States.
The solubility of commercial dry alum at various temperatures is listed in Table 10-2:
TABLE 10-2
SOLUBILITY OF ALUM AT VARIOUS TEMPERATURES
Temperature Solubility
°F Ib/gal
32 6.03
50 6.56
68 7.28
86 8.45
104 10.16
Dry alum is not corrosive unless it absorbs moisture from the air, such as during prolonged
exposure to humid atmospheres. Therefore, precautions should be taken to ensure storage
space is free of moisture.
Alum is shipped in 100 Ib bags, drums, or in bulk (minimum of 40,000 Ib) by truck
or rail. Bag shipments may be ordered on wood pallets. In 1973, the price range for
dry alum in bulk quantities is $60 to $65 per ton, F.O.B. the point of manufacture.
Freight costs must be added to this. Bagging adds about $5 per ton to the base cost
for bulk alum. Names and locations of major producers are in a subsequent section.
10.1.1.2 General Design Considerations
Water utilities usually use ground and rice alum because of their superior flow
characteristics. These grades have less tendency to lump or arch in storage and therefore
provide more consistent feeding qualities. Hopper agitation is seldom required with these
grades, and in fact may be determined to feeding due to packing the contents of the
bin.
10-2
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Alum dust is present in the ground grade and will cause minor irritation of eyes and
respiratory tract. A respirator will protect against alum dust. Alum dust should be
thoroughly flushed from the eyes immediately and washed from the skin with water.
Gloves should be worn to protect the hands. Because of minor irritation in handling
and the possibility of alum dust causing rusting of adjacent machinery, dust removal
equipment is desirable.
10.1.1.3 Storage
A typical storage and feeding system for dry alum is shown in Figure 10-1. Bulk alum
can be stored in mild steel or concrete bins with dust collector vents located in, above,
or adjacent to the equipment room. Recommended storage capacity is about 30 days.
Dry alum in bulk can be transferred with screw conveyors, pneumatic conveyors, or bucket
elevators made of mild steel. Pneumatic conveyor elbows should have a reinforced backing
to withstand abrasion.
Bags and drums of alum should be stored in a dry location. Bag or drum loaded hoppers
should have storage capacity for eight hours at the nominal maximum feed rate so personnel
are not required to charge the hopper more than once per shift. Converging hopper sections
should have a minimum wall slope of 60° to prevent arching.
Bulk Storage hoppers should have a discharge bin gate so feeding equipment may be isolated
for servicing. The bin gate should be followed by a flexible connection, and transition
hopper chute or hopper which acts as a conditioning chamber over the feeder.
10.1.1.4 Feeding Equipment
The feed system includes all components required for preparation of the chemical solution.
Capacities and assemblies should be selected to fulfill individual system requirements. Three
basic types of chemical feed equipment are used: volumetric, belt gravimetric, and
loss-in-weight gravimetric. Volumetric feeders are usually used where low initial cost and
lower delivery capacities are the basis of selection. Volumetric feeder mechanisms are
usually exposed to corrosive vapors from the dissolving chamber. Manufacturers usually
control this problem by use of an electric heater to keep the feeder housing dry or by
using plastic components in the exposed areas.
Volumetric dry feeders are generally of the screw type. Two designs of screw feed
mechanism are available. Both allow even withdrawal across the bottom of the feeder
hopper to prevent dead zones. One screw design is the variable pitch type with the pitch
expanding evenly to the discharge point. The second screw design is the constant
pitch-reciprocating type. This type has each half of the screw turned in opposite directions
so the turning and reciprocating motion alternately fills one half of the screw while the
other half is discharging. The variable pitch screw has one point of discharge, while the
constant pitch-reciprocating screw has two points of discharge, one at each end of the
10-3
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FIGURE 10-1
TYPICAL DRY FEED SYSTEM
,-DUST COLLECTOR
/>
XX-FILL PIPE (PNEUMATIC)
K_
BULK STORAGE
BIN
DAY HOPPER
FOR DRY CHEMICAL
FROM BAGS OR DRUMS,
*-<•• •'••••;.-
BIN GATE
FLEXIBLE
CONNECTION
ALTERNATE SUPPLIES DEPENDING
ON STORAGE
DUST COLLECTOR
XBAG FILL
-SCREEN
WITH BREAKER
SCALE OR SAMPLE CHUTE
DUST AND VAPOR REMOVER
WATER
SUPPLY
DRAIN
SOLENOID VALVE
CONTROL
VALVE -^ROTAMETER
PRESSURE REDUCING
VALVE
I
WATER SUPPLY
__ GRAVITY TO
APPLICATION
PUMP
APPLICATION
10-4
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screw. The accuracy of volumetric feeders is influenced by the character of material
being fed and ranges between 1% for free-flowing materials to 7% for cohesive materials.
This accuracy is volumetric and is not related to accuracy by weight (gravimetric).
Where the greatest accuracy and more economical use of chemical is desired, the
loss-in-weight type feeder should be selected. This feeder is suited to low and medium
feed rates with a maximum of approximately 4,000 Ib/hour. The unit consists of a material
hopper and feeding mechanism mounted on enclosed scales. The feed rate controller
retracts scale poise weight to deliver the dry chemical at the desired rate. The feeding
mechanism must feed at this rate to maintain the balance of the scale. Any unbalance
of the scale beam causes a corrective change in output. Continuous comparison of actual
hopper weight with set hopper weight prevents cumulative errors. Accuracy of the
loss-in-weight feeder is near 1% by weight of the set rate.
Belt-type gravimetric feeders span the capacity ranges of volumetric and loss-in-weight
feeders and can be sized for most applications in wastewater treatment. Initial expense
falls between that for the volumetric feeder and the loss-in-weight feeder. Belt-type
gravimetric feeders consist of a basic belt feeder incorporating a weighing and control
system. Feed rates can be varied by changing either the weight per foot of belt, belt
speed, or both. Controllers in general use are mechanical, pneumatic, electric and
mechanical-vibrating. Accuracy specified for belt-type gravimetric feeders should be within
1% of set rate. This equipment normally includes mild steel hoppers, stainless steel
mechanism components, and rubber surfaced feed belts.
Because alum solution is corrosive, dissolving or solution chambers should be constructed
of type 316 stainless steel, fiberglass reinforced plastic (FRP), and plastics. Dissolvers
should be sized for preparation of the desired solution strength. The most dilute solution
strength usually recommended is 0.5 Ib of alum to 1 gallon of water, or a 6% solution.
The dissolving chamber is designed for a minimum detention time of 5 minutes at the
maximum feed rate. Because excessive dilution may be detrimental to coagulation,
eductors, or float valves that would ordinarily be used ahead of centrifugal pumps, are
not recommended. Dissolvers should be equipped with water meters and mechanical mixers
so the water to alum ratio may be properly established and controlled.
10.1.1.5 Piping and Accessories
Pipe made from FRP or plastics (polyvinyl chloride, polyethylene, polypropylene, and
similar materials) is recommended for alum solution. Care must be taken to provide
adequate support for these piping systems, with close attention given to spans between
supports so objectionable deflection will not occur. Lined steel pipe is generally tougher
and more rigid, but the cost of providing a near perfect lining may detract from its
suitability.
10-5
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Solution flow by gravity to the point of discharge is desirable. When gravity flow is
not possible, transfer components should require little or no dilution. When metering
pumps or proportioning weir tanks are used, return of excess flow to a holding tank
should be considered. Metering pumps are discussed in the section on liquid alum.
Valves used in solution lines should be plastic, type 316 stainless steel or rubber-lined
iron or steel.
10.1.1.6 Pacing and Control
Volumetric and gravimetric feeders are usually adaptable to operation from any standard
instrument control and pacing signals.
When solution must be pumped, consideration should be given to use of holding tanks
between the dry feed system and feed pumps; solution water supply should be controlled
to prevent excessive dilution. The dry feeders may be started and stopped from tank
level probes. Variable control metering pumps can transfer alum stock solution to the
point of application without further dilution.
Means should be provided for calibration of chemical feeders. Volumetric feeders may
be mounted on platform scales. Belt feeders should include a sample chute and box
to check actual delivery with set delivery.
Gravimetric feeders are usually furnished with totalizers only. Remote instrumentation
is frequently used with gravimetric equipment, but seldom used with volumetric equipment.
10.1.2 Liquid Alum
10.1.2.1 Properties and Availability
Liquid alum is shipped in insulated tank cars or trucks. During the winter it is heated
prior to shipment so crystallization will not occur during transit. Liquid alum is shipped
at a solution strength of about 4.37% as aluminum or about 8.3% as A^O^ or 49% as
Al2(SO4)3 14H2O. This solution has a density of 11.1 Ib/gal at 60°F and contains
about 5.4 Ib dry alum (17% A^Oj) per gallon of liquid. This solution will begin to
crystallize at 30°F and crystallizes at 18°F.
10-6
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Crystallization temperatures of other solution strengths are as follows (1) in Table 10-3:
TABLE 10-3
CRYSTALLIZATION TEMPERATURES OF LIQUID ALUM
Temperature of
Crystallization
op
5.19 26
6.42 21
6.67 19
6.91 17
7.16 15
7.40 13
7.66 12
7.92 14
8.19 17
8.46 20
8.74 28
The viscosity of various alum solutions is given in Figure 10-2.
Tank truck lots of 3,000 to 5,000 gallons are available. Tank car lots are available in
quantities of 7,000 to 18,000 gallons. The 1974 price of liquid alum on an equivalent
dry alum (17% A^O^) basis is about $60 per ton, F.O.B. the point of manufacture.
Since liquid alum is an intermediate compound in production of dry alum, the liquid
form costs less. However, liquid alum costs more to transport since it is nearly half
water, by weight. Therefore, a cost tradeoff point can be established when exact chemical
costs and local freight rates are determined. Liquid alum will generally be more economical
than dry alum if the point of use is within 100 miles of the manufacturing plant; ease
of handling, storing and feeding liquid alum extend its practical transport limit to 200 miles
or more.
10-7
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CO
LLl
V)
UJ
o
CO
o
o
V)
FIGURE 10-2
VISCOSITY OF ALUM SOLUTIONS
(COURTESY OF ALLIED CHEMICAL CO.)
I I I I I I I I I I 1 I I I I I I
30 40 50 60 70 80 90 100 110 120 130 140 150 160170 180 190 200 210 220
TEMPERATURE,°F
10-8
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10.1.2.2 General Design Considerations
Bulk unloading facilities usually must be provided at the treatment plant. Rail cars are
constructed for top unloading and therefore require an air supply system and flexible
connectors to pneumatically displace the alum from the car. U.S. Department of
Transportation regulations concerning chemical tank car unloading should be observed.
Tank truck unloading is usually accomplished by gravity or by a truck mounted pump.
Established practice in the treatment field has been to dilute liquid alum prior to
application. However, recent studies have shown that feeding undiluted liquid alum results
in better coagulation and settling. This is reportedly due to prevention of hydrolysis
of the alum.
No particular industrial hazards are encountered in handling liquid alum. However, a
face shield and gloves should be worn around leaking equipment. The eyes or skin should
be flushed and washed upon contact with liquid alum. Liquid alum becomes very slick
upon evaporation and therefore spillage should be avoided.
10.1.2.3 Storage
Liquid alum is stored without dilution at the shipping concentration. Storage tanks may
be open if indoors but must be closed and vented if outdoors. Outdoor tanks should
also be heated, if necessary, to keep the temperature above 25°F to prevent crystallization.
Storage tanks should be constructed of type 316 stainless steel; FRP, steel lined with
rubber, polyvinyl chloride, or lead; see subsequent section for further details. Liquid
alum can be stored indefinitely without deterioration.
Storage tanks should be sized according to maximum feed rate, shipping time required,
and quantity of shipment. Tanks should generally be sized for 1-1/2 times the quantity
of shipments. A 10-day to 2-week supply should be provided to allow for unforeseen
shipping delays.
10.1.2.4 Feeding Equipment
Various types of gravity or pressure feeding and metering units are available. Figures 10-3
and 10-4 illustrate commonly used feed systems. The rotodip-type feeder or rotameter
is often used for gravity feed and the metering pump for pressure feed systems.
The pressure or head available at the point of application frequently determines the feeding
system to be used. The rotodip feeder can be supplied from overhead storage by gravity
with the use of an internal level control valve, as shown by Figure 10-3. It may also
10-9
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FIGURE 10-3
ALTERNATIVE LIQUID FEED SYSTEMS
FOR OVERHEAD STORAGE
r-
FLOAT
VALVE-v
®-rf
7^
ROTOD IP-TYPE FEEDER
OVERHEAD
STORAGE
TANK
CONTROL
VALVE
*
ROTAMETER
UJ
Ul
ce
1
O
F VALVE
K PRESSURE VALVE
r
METERING PUMP
GRAVITY FEED
10-10
-------
FIGURE 10-4
ALTERNATIVE LIQUID FEED SYSTEMS
FOR GROUND STORAGE
PRESSURE FEED
/
(
GRAVITY
,-ROTOD IP-TYPE
/CTx r
\SJ\7 *
.
7
FEED
TRANSFER PUMF
•
FEEDER
i
'
\
^CONTROL VALVE
y ^ROTAMETER
n Jf 1. 1
^ 1 -
s n tv
^
IRCULATION
GROUND
STORAGE
TANK
1
•>• j
.j Q- iij
•< a. Z oe
§ i E i ^
u. a. oe co
_ CD LU
oe ci z ae
Z UJ OS ' ' °"
0 uj "• S
•"s !s AS
Q T f f
1 0
PRESSURE
FEED
1 .
>,
10-11
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be supplied by a centrifugal pump. The latter arrangement requires an excess flow return
line to the storage tank, as shown by Figure 10-4. Centrifugal pumps should be
direct-connected but not close-coupled because of possible leakage into the motor, and
should be constructed of type 316 stainless steel, FRP, and plastics.
Metering pumps, currently available, allow a wide range of capacity to compare with the
rotodip and rotameter systems. Hydraulic diaphragm type pumps are preferable to other
type pumps and should be protected with an internal or external relief valve. A back
pressure valve is usually required in the pump discharge to provide efficient check valve
action. Materials of construction for feeding equipment should be as recommended by
the manufacturer for the service, but depending on the type of system, will generally
include type 316 stainless steel, FRP, plastics, and rubber.
10.2.1.5 Piping and Accessories
Piping systems for alum should be FRP, plastics (subject to temperature limits), type 316
stainless steel, or lead. Piping and valves used for alum solutions are also discussed in
the preceding section on dry alum.
10.1.2.6 Pacing and Control
The feeding systems described above are volumetric, and the feeders generally available
can be adapted to receive standard instrument pacing signals. The signals can be used
to vary motor speed, variable speed transmission setting, stroke speed and stroke length
where applicable. A totalizer is usually furnished with a rotodip-type feeder, and remote
instruments are available. Instrumentation is rarely used with rotameters and metering
pumps.
10.1.3 Dry Sodium Aluminate
10.1.3.1 Properties and Availability
Dry sodium aluminate, Na2Al2C>4, is available from manufacturers listed in a subsequent
section. Dry sodium aluminate is shipped in 50 Ib bags and has a bulk density ranging
from 40 to 50 Ib/ft . The A^Og content ranges from 41 to 46%. Dry sodium aluminate
is noncorrosive and the pH of a 1% solution is about 11.9. Manufacturers should be
consulted for more precise specifications of their product.
The 1974 price of dry sodium aluminate in 40,000 Ib shipments ranges from $0.13 to
$0.15 per Ib, F.O.B. the point of manufacture. Prices increase substantially for smaller
shipments.
10-12
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10.1.3.2 General Design Considerations
Requirements for dry sodium aluminate feed systems are generally similar to those for
dry aluminum sulfate. Dry sodium aluminate is not available in bulk quantities. Therefore,
the small, day-type hoppers with manual filling arrangements as shown by Figure 10-1
are used. Precautionary measures for handling sodium aluminate are similar to those for
strong alkalies, such as caustic soda. Contact with skin, eyes, and clothing should be
avoided. Aluminate dust or solution spray should not be breathed.
10.1.3.3 Storage
Dry sodium aluminate is stored as received, in bags, and at optimum conditions of 60
to 90°F; the recommended storage limit is 6 months. Hopper material of mild steel
is completely adequate. This chemical may or may not be free flowing, depending on
the manufacturer and grade used. Therefore, hopper agitation may be required. Sodium
aluminate deteriorates on exposure to the atmosphere and care should be taken to avoid
tearing of bags.
10.1.3.4 Feeding Equipment
Materials of construction for dissolving chambers may be mild steel or stainless steel and
selection may be influenced by conformity with adjacent equipment. Equipment similar
to that shown by Figure 10-1 is applicable. Standard practice for the free-flowing grade
of sodium aluminate calls for dissolvers sized for 0.5 Ib per gallon or 6% solution strength
with a dissolver detention time of 5 minutes at the maximum feed rate.
After dissolving dry sodium aluminate in the preparation of batch solutions, agitation
should be minimized or eliminated to prevent deterioration of the solution. Air agitation
is not recommended, and solution tanks should be covered to prevent carbonation of
the solution.
10.1.3.5 Piping and Accessories
Materials for piping and transporting dry sodium aluminate solution may be mild steel,
iron, type 304 stainless steel, concrete, or plastics. The use of copper, copper alloys,
and rubber should be avoided.
10-13
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10.1.3.6 Pacing and Control
Pacing and control fundamentals are similar to those described for dry alum. The amount
of dilution is not a consideration in the use of sodium aluminate. Therefore, the use
of float valves to satisfy centrifugal pump suction, and the use of eductors, is permissible.
10.1.4 Liquid Sodium Aluminate
10.1.4.1 Properties and Availability
Liquid sodium aluminate is available from the manufacturers listed in a subsequent section.
There is considerable variety in the composition of liquid sodium aluminate from the
manufacturers listed. The A^Og content varies from 4.9 to 26.7%. The lower solution
strengths are usually more expensive because of the cost of freighting the solution water.
Because of the variety of solution strengths available, the manufacturers should be
contacted for more specific information on density, viscosity, and cost.
Liquid sodium aluminate is available in 30-gallon drums (380 Ib), tank truck, and tank
car quantities. The approximate 1974 price of liquid sodium aluminate in 40,000 Ib
quantities is about $0.05 per Ib (A12O3 content, 26.4% or 19.9%) P.O. B. point of
manufacture.
10.1.4.2 General Design Considerations
Because of the alkaline nature of sodium aluminate, it should not be used in contact
with brass, copper, aluminum or rubber. Liquid sodium aluminate is a strong alkali and
the same precautions should be exercised in handling it as in handling caustic soda.
10.1.4.3 Storage
Liquid sodium aluminate is usually stored at the shipping concentration, either in the
shipping drums or in mild steel tanks. Storage tanks may be located indoors or outdoors;
however, outdoor tanks should be provided with facilities for indirect heating. The
maximum recommended length of storage is two to three months. Bulk shipments can
be unloaded by gravity, pumping, or air pressure. However, if air is used, it should first
be passed through lime-caustic soda breathers to remove carbon dioxide. Steam injection
facilities are required at the unloading site.
10.1.4.4 Feeding Equipment
Feeding equipment and systems as described for liquid alum generally apply to liquid
sodium aluminate except with changes of requirements regarding dilution and materials
of construction as described above.
10-14
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Liquid sodium aluminate may be fed at shipping strength or diluted to a stable 5 to
10% solution. Stable solutions are prepared by direct addition of low hardness water
and mild agitation. Air agitation is not recommended.
10.1.4.5 Piping and Accessories
Piping requirements are the same as previously indicated for solutions of dry aluminate.
10.1.4.6 Pacing and Control
System pacing and control requirements are the same as described for liquid alum.
10.1.5 Aluminum Chloride
Aluminum chloride is not as readily available as other compounds previously discussed,
and in some cases it does not enjoy a competitive price structure; nevertheless, it should
be considered as a source of aluminum by designers undertaking phosphorus precipitation
by metal addition.
The majority of the aluminum chloride production facilities in this country are located
near petroleum refining centers and petrochemical complexes. This is because those
industries consume three-fourths of the total production. Miscellaneous users of the balance
of the overall production include makers of antiperspirants and a variety of other unrelated
manufacturing operations. Therefore, a reliable and economical source of aluminum
chloride may, in fact, be near the proposed wastewater treatment project.
Much of the aluminum chloride produced is the form of semi-pure anhydrous crystals
with an off-white color. This product is derived from direct chlorination of scrap
aluminum, and has a molecular weight of 133.3 and purities near 99 percent. In some
instances purities may drop to 96 percent which is satisfactory for wastewater treatment
if analysis shows other compounds present are acceptable.
Another solid form is produced by crystallization from hydrochloric acid in the form
of a hexahydrate, in which six molecules of water attach to each molecule of aluminum
chloride.
The most usual commercial form of liquid aluminum chloride is in a 32-degree Baume
solution; it contains 28 percent aluminum chloride by weight. This particular liquor is
not universally produced, however, so designers should be prepared to consider other
concentrations with a fairly wide variety of other chemical compounds present.
10-15
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Containers for shipping and storing should be similar to those used for dry and liquid
alum. Handling and feeding requirements are also similar to those used with alum.
Designers should check with local producers for further details.
Designers may contact manufacturers operating in at least seventeen plants in the U.S.
shown in Table 10-4:
TABLE 10-4
SUPPLIERS OF ALUMINUM CHLORIDE
Allied Chemical Corp.:
Clinton Chemical Co.:
Dow Chemical Co.:
E. I. du Pont de Nemours
& Co., Inc.:
FMC Corporation:
Hercules Alchlor
Chemical Co., Inc.
Monsanto Co.:
Pearsall Chemical Corp.
Reheis Co., Inc.:
Stauffer Chemical Co.:
El Segundo, Calif.;
Hegewisch, 111.;
Elberta, N.Y.
Houston, Texas
Freeport, Texas
Grasselli, N.J.;
East Chicago, Ind.
Nitro, W. Va.
Ravenna, Ohio
Everett, Mass.;
Texas City, Texas
Phillipsburg, N.J.;
La Porte, Texas
Berkeley Heights, N.J.
Dominguez, Calif.;
Baton Rouge, La.;
Elkton, Md.
10-16
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10.1.6 Estimated Initial Costs of Adding Liquid Alum Feed Facilities
to Existing Wastewater Treatment Plants
Cost estimates of chemical storage and feeding equipment for 1, 10, and 100 mgd plants
have been prepared based on the use of liquid alum (8.3% A^O^) for phosphorus
precipitation. A mole ratio of A1:P of 2:1 was assumed to treat an influent phosphorus
concentration of 10 mg/1 as P. This corresponds to an alum dose of 191 mg/1 (17.4 mg/1
as aluminum).
Chemical feed equipment was sized for a peak feed rate of twice that calculated from
the mole ratio. Storage was provided for at least 15 days at the average feed rate. Th'*-
storage time is arbitrary and will vary at each installation depending on the distance to
and the reliability of the source of chemical supply. Piping and buildings to house the
feeding equipment are not included in the estimates. At many locations the chemical
feeding equipment can be installed in existing buildings. In all estimates, the costs include
an allowance for the contractor's installation, overhead, and profit, and an allowance of
20% of the construction cost for engineering and contingencies. The estimated costs for
each size plant are shown in Table 10-5. These figures correspond to an EPA Construction
Cost Index of 185:
TABLE 10-5
COST OF LIQUID ALUM FEED FACILITIES
Plant Size Estimated 1973 Cost
(mgd) ($)
1 15,000
10 50,000
100 600,000
For the 1.0 mgd plant, feeding equipment for liquid alum includes two 25 gph hydraulic
diaphragm pumps (one operating and one standby) with the necessary accessories and
equipment to pace the feed rate with plant flow. Two 3,000-gallon FRP tanks with
accessories are provided for storage of liquid alum. The total capacity of 6,000 gallons
allows some flexibility in ordering and receiving 4,000-gallon tank truck shipments.
10-17
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For the 10 mgd plant, liquid alum is fed by three 125 gph rotodip-type feeders (two
operating and one standby) with the necessary accessories and control panel for
proportioning chemical feed to flow. Alum is stored in four 11,500-gallon FRP tanks.
For the 100 mgd plant, liquid alum is fed with seven 415 gph rotodip-type feeders (six
operating and one standby) with the necessary control equipment. Storage costs are based
on ten 50,000-gallon underground concrete tanks for the rotodip feeders but eliminates
the need for heating the tanks. Three 500-gallon FRP day tanks are included in the
estimate; one for each pair of feeders. Four alum transfer pumps (three operating and
one standby), with a capacity of 50 gpm at 50 ft of head are provided.
The cost of feeding and storage facilities for the 100 mgd plant is greater than 10 times
the cost of facilities for the 10 mgd plant. This is because underground storage facilities
were the basis of design for the 100 mgd plant in contrast to FRP tanks located in an
existing building for the 10 mgd plant. The cost of a building for the 1.0 and 10.0 mgd
plants would raise the cost substantially. FRP tanks could be used for the larger plant;
however, structural design may necessitate special construction requirements for
underground FRP tanks. FRP tanks located above ground and on concrete foundations
would require special insulation and heating.
10.2 Iron Compounds
10.2.1 Liquid Ferric Chloride
10.2.1.1 Properties and Availability
Liquid ferric chloride (2) is a corrosive, dark brown oily-appearing solution having a typical
unit weight between 11.2 and 12.4 Ib/gal (35 to 45% FeCl3). The ferric chloride content
of these solutions is 3.95 to 5.58 Ib/gal. Shipping concentrations vary from summer to
winter due to the relatively high crystallization temperature of the more concentrated
solutions as shown by Figure 10-5. The pH of a 1% solution is 2.0.
The molecular weight of ferric chloride is 162.22. Viscosities of ferric chloride solutions
at various temperatures are presented in Figure 10-6.
Liquid ferric chloride is shipped in 3,000 to 4,000 gallon bulk truckload lots, in 4,000
to 10,000 gallon bulk carload lots, and in 5 and 13 gallon carboys. Liquid ferric chloride
is produced at the locations listed in a subsequent section. The 1974 price of liquid
ferric chloride in bulk quantities is $0.045 per Ib (as ferric chloride), F.O.B. point of
manufacture.
10-18
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FIGURE 10-5
FREEZING POINT CURVE FOR COMMERCIAL
FERRIC CHLORIDE SOLUTIONS
(ADAPTED FROF1 DOW CHEMICAL CO.)
UJ
tr
UJ
a.
UJ
i 1 1 1
WATER UTILITY GRADE
-20
3O 40
PERCENT BY WEIGHT IN WATER
10-19
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8
100
80
60
50
40
30
20
15
10
8
6
5
4
FIGURE 10-6
VISCOSITY VS COMPOSITION OF FERRIC CHLORIDE
SOLUTIONS AT VARIOUS TEMPERATURES
(COURTESY OF DOW CHEMICAL CO.)
I I I I I I I I I I
(Absolute Viscositjr)=(Kinematic Viscosity)(Density)
Cent!poises = Cent!stokes x Gin
cc
10
20
50
10-20
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Tank trucks and cars are normally unloaded pneumatically, and operating procedures must
be closely followed to avoid spills and accidents. The safety vent cap and assembly (painted
red) should be removed prior to opening the unloading connection to depressurize the
tank car or truck, prior to unloading.
10.2.1.2 General Design Considerations
Ferric chloride solutions are corrosive to many common materials and cause stains which
are difficult to remove. Areas which are subject to staining should be protected with
resistant paint or rubber mats.
Normal precautions should be employed when cleaning ferric chloride handling equipment.
Workmen should wear rubber gloves, rubber apron, and goggles or a face shield. If ferric
chloride comes in contact with the eyes or skin, flush with copious quantities of running
water and call a physician. If ferric chloride is ingested, induce vomiting and call a
physician.
10.2.1.3 Storage
Ferric chloride solution can be stored as shipped. Storage tanks should have a free vent
or vacuum relief valve. Tanks may be constructed of FRP, rubber lined steel, or plastic
lined steel. Resin impregnated carbon or graphite are also suitable materials for storage
containers.
It may be necessary in most instances to house liquid ferric chloride tanks in heated
areas or provide tank heaters or insulation to prevent crystallization. Ferric chloride can
be stored for long periods of time without deterioration. The total storage capacity should
be 1-1/2 times the largest anticipated shipment, and should provide at least a 10-day to
2-week supply of the chemical at the design average dose.
10.2.1.4 Feeding Equipment
Feeding equipment and systems described for liquid alum generally apply to ferric chloride
except for materials of construction, and the use of glass tube rotameters.
It may not be desirable to dilute the ferric chloride solution from its shipping concentration
to a weaker feed solution because of possible hydrolysis. Ferric chloride solutions may
be transferred from underground storage to day tanks with impervious graphite or rubber
lined self-priming centrifugal pumps having teflon rotary and stationary seals. Because
10-21
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of the tendency for liquid ferric chloride to stain or deposit, glass-tube rotameters should
not be used for metering this solution. Rotodip feeders and diaphragm metering pumps
are often used for ferric chloride, and should be constructed of materials such as rubber
lined steel and plastics.
10.2.1.5 Piping and Accessories
Materials for piping and transporting ferric chloride should be rubber or Saran lined steel,
hard rubber, FRP, or plastics. Varying should consist of rubber or resin lined diaphragm
valves, Saran lined valves with teflon diaphragms, rubber sleeved pinch type valves, or
plastic ball valves. Gasket material for large openings such as manholes in storage tanks
should be soft rubber; all other gaskets should be graphite-impregnated blue asbestos, teflon,
or vinyl.
10.2.1.6 Pacing and Control
System pacing and control requirements are similar to those discussed previously for liquid
alum.
10.2.2 Ferrous Chloride (Waste Pickle Liquor)
10.2.2.1 Properties and Availability
Ferrous chloride, FeC^, as a liquid is available in the form of waste pickle liquor from
steel processing. The liquor weighs between 9.9 and 10.4 Ib/gal and contains 20 to 25%
ferrous chloride or about 10% available iron. A 22% solution of ferrous chloride will
crystallize at a temperature of -4°F. The molecular weight of ferrous chloride is 126.76.
Free acid in waste pickle liquor can vary from 1 to 10% and usually averages about 1.5
to 1.0%. Ferrous chloride is slightly less corrosive than ferric chloride.
Waste pickle liquor is available in 4,000-gallon truckload lots and a variety of carload
lots. In most instances the availability of waste pickle liquor will depend on the proximity
to steel processing plants. Dow Chemical Company produces a waste pickle liquor, having
a ferrous chloride content of about 22% at a 1974 price of $0.045 per Ib of ferrous
chloride in bulk car or truckload quantities, F.O.B. Midland, Michigan.
10.2.2.2 General Design Considerations
Since ferrous chloride or waste pickle liquor may not be available on a continuous basis,
storage and feeding equipment should be suitable for handling ferric chloride. Therefore,
the ferric chloride section should be referred to for storage and handling details.
10-22
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10.2.3 Ferric Sulfate
10.2.3.1 Properties and Availability
Ferric sulfate, Fe2(SO4)3 • X H2O, is marketed as a dry, partially hydrated product with
seven water molecules. Typical properties are presented in Table 10-6 below (3):
TABLE 10-6
PROPERTIES OF FERRIC SULFATE
Molecular Weight 526
Bulk Density, Ib/cu ft 56-60
Water Soluble Iron
Expressed as Fe 21.5%
Water Soluble Fe+3 19.5%
Insolubles Total 2.0%
Free Acid 2.5%
Moisture @ 105°C 2.0%
Ferric sulfate is shipped in car and truckload lots of 50 Ib and 100 Ib moistureproof
paper bags and 200 Ib and 400 Ib fiber drums. Bulk carload shipments in box and closed
hopper cars are available. The major producer is Cities Service Company, with a plant
located at Copper Hill, Tennessee.
The 1973 price of bulk ferric sulfate (21.8% Fe) is about $39 per ton, F.O.B. Copper
Hill, Tennessee. Bagging adds some $6 to $11 per ton over the bulk rate.
General precautions should be observed when handling ferric sulfate, such as wearing goggles
and dust masks, and areas of the body that come in contact with the dust or vapor
should be washed promptly.
10-23
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10.2.3.2 General Design Considerations
Aeration of ferric sulfate should be held to a minimum because of the hygroscopic nature
of the material, particularly in damp atmospheres. Mixing of ferric sulfate and quicklime
in conveying and dust vent systems should be avoided as caking and excessive heating
can result. The presence of ferric sulfate and lime in combination has been known to
destroy cloth bags in pneumatic unloading devices (4). Because ferric sulfate in the presence
of moisture will stain, precautions similar to those discussed for ferric chloride should
be observed.
10.2.3.3 Storage
Ferric sulfate is usually stored in the dry state either in the shipping bags or in bulk
in concrete or steel bins. Bulk storage bins should be as tight as possible to avoid moisture
absorption, but dust collector vents are permissible and desirable. Hopper walls should
slope 36% or more.
Bins may be located inside or outside and the material transferred by bucket elevator,
screw or air conveyors. Ferric sulfate stored in bins usually absorbs some moisture and
forms a thin protective crust which retards further absorption until the crust is broken.
10.2.3.4 Feeding Equipment
Feed solutions are usually made up at a water to chemical ratio of 2:1 to 8:1 (on a
weight basis) with the usual ratio being 4:1 with a 20-minute detention time. Care must
be taken not to dilute ferric sulfate solutions to less than 1% to prevent hydrolysis and
deposition of ferric hydroxide. Ferric sulfate is actively corrosive in solution, and dissolving
and transporting equipment should be fabricated of type 316 stainless steel, rubber, plastics,
ceramics or lead.
Dry feeding requirements are similar to those for dry alum except that belt type feeders
are rarely used because of their open type of construction. Closed construction, as found
in the volumetric and loss-in-weight type feeders, generally exposes a minimum of operating
components to the vapor, and thereby minimizes maintenance. A water jet vapor remover
should be provided at the dissolver to protect both the machinery and operator.
10.2.3.5 Piping and Accessories
Piping systems for ferric sulfate should be FRP, plastics, type 316 stainless steel, rubber
or glass.
10.2.3.6 Pacing and Control
System pacing and control are the same as discussed for dry alum.
10-24
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10.2.4 Ferrous Sulfate
10.2.4.1 Properties and Availability
Ferrous sulfate or copperas is a byproduct of pickling steel and is produced as granules,
crystals, powder, and lumps. The most common commercial form of ferrous sulfate is
FeSO4 • 7H2O, with a molecular weight of 278, and containing 55 to 58% FeSO4 and
20 to 21% Fe. The product has a bulk density of 62 to 66 lb/ft3. When dissolved,
ferrous sulfate is acidic. The composition of ferrous sulfate may be quite variable and
should be established by consulting the nearest manufacturers.
Bulk, drum (400-lb) and bag (50 and 100 Ib) shipments are available from producers
at locations listed in a subsequent section. The 1974 bulk price of ferrous sulfate in
bulk carload and truckload quantities is about $28 per ton (21% Fe), with an additional
$8 per ton for bagging.
Ferrous sulfate is also available in a wet state in bulk form from some plants. This form
is likely to be difficult to handle and the manufacturer should be consulted for specific
information and instructions.
Dry ferrous sulfate cakes at storage temperatures above 68°F, is efflorescent in dry air,
and oxidizes and hydrates further in moist air.
General precautions similar to those for ferric sulfate with respect to dust and handling
acidic solutions, should be observed when working with ferrous sulfate. Mixing quicklime
and ferrous sulfate produces high temperatures and the possibility of fire.
10.2.4.2 General Design Considerations
The granular form of ferrous sulfate has the best feeding characteristics and gravimetric
or volumetric feeding equipment may be used.
The optimum chemical to water ratio for continuous dissolving is 0.5 Ib/gal or 6% with
a detention time of 5 minutes in the dissolvers. Mechanical agitation should be provided
in the dissolver to assure complete solution. Lead, rubber, iron, plastics, and type 304
stainless steel can be used as construction materials for handling solutions of ferrous sulfate.
Storage, feeding and transporting systems probably should be suitable for handling ferric
sulfate as an alternative to ferrous sulfate.
10.2.5 Estimated Initial Costs of Adding Liquid Ferric Chloride Feed
Systems to Existing Wastewater Treatment Plants
10-25
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Costs of chemical storage and feeding equipment for 1, 10, and 100 mgd plants have
been estimated based on the use of liquid ferric chloride (basis: 35% ferric chloride)
for phosphorus precipitation. A mole ratio of Fe:P of 2.0:1 was assumed to treat an
influent phosphorus concentration of 10 mg/1 as P. This corresponds to a ferric chloride
dose of 104 mg/1 (36 mg/1 as Fe3+).
Chemical feed equipment was sized for a peak feed rate of twice that calculated from
the mole ratio. Storage was provided for at least 15 days at the average feed rate. This
storage time is arbitrary and will vary at each installation depending on the distance to
and the reliability of the souce of chemical supply. Piping and buildings to house the
feeding equipment are not included. At many locations the chemical feeding equipment
can be installed in existing buildings. All estimated costs include an allowance for the
contractor's installation, overhead, and profit plus an allowance of 20% of the construction
cost for engineering and contingencies. The estimated costs for each size plant are shown
in Table 10-7. These figures correspond to an EPA Construction Cost Index of 185:
TABLE 10-7
COSTS OF LIQUID FERRIC CHLORIDE FEED FACILITIES
Plant Size Estimated 1973 Cost
(mgd) ($)
1 15,000
10 50,000
100 500,000
For the 1 mgd plant, the feeding equipment for liquid ferric chloride includes two 18 gph,
hydraulic diaphragm pumps (one operating and one standby) with the necessary accessories
and equipment to pace the feed rate to the plant flow. Liquid ferric chloride is stored
in two 3,000-gallon FRP tanks with accessories. The 6,000-gallon total capacity allows
use of liquid ferric chloride in 4,000-gallon tank truck quantities.
For the 10 mgd plant, liquid ferric chloride is fed by three 90 gph rotodip-type feeders
(two operating and one standby) with the necessary accessories and control panel for
proportioning chemical feed to flow.
For the 100 mgd plant, seven 305 gph rotodip-type feeders (six operating and one standby)
are provided along with the necessary control equipment. Storage is provided in eight
50,000-gallon rubber-lined concrete tanks located underground. The underground storage
necessitates transfer pumps and day tanks for the rotodip feeders, but eliminates the need
10-26
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for heating the tanks. Three 500-gallon FRP day tanks are included along with four
ferric chloride transfer pumps sized for 35 gpm each at 50 ft of head.
With regard to the cost of the feed system for the 100 mgd plant being out of proportion
to that for a 10 mgd plant, the reader is referred to Section 10.1.5 of this chapter for
the discussion of underground concrete storage vessels as opposed to FRP tanks.
10.3 Lime
10.3.1 Quicklime
10.3.1.1 Properties and Availability
Quicklime, CaO, has a density range of approximately 55 to 75 lb/ft3, and a molecular
weight of 56.08. A slurry for feeding, called milk of lime, can be prepared with up
to 45% solids. Lime is only slightly soluble, and both lime dust and slurries are caustic
in nature. A saturated solution of lime has a pH of about 12.4.
Lime can be purchased in bulk in both car and truckload lots. It is also shipped in
80 and 100 Ib multiwall "moistureproof" paper bags. Lime is produced at the locations
listed in a subsequent section (5).
1974 prices for bulk pebble quicklime range from $18 per ton to $22 per ton with the
higher prices generally in the far west, and higher than average in the north. Bagging
adds approximately $5 per ton to the cost.
The CaO content of commercially available quicklime can vary quite widely over an
approximately range of 70 to 96%. Content below 88% is generally considered below
standard in the municipal use field (6). Purchase contracts are often based on 90% CaO
content with provisions for payment of a bonus for each 1% over and a penalty for each
1% under the standard. A CaO content less than 75% probably should be rejected because
of excessive grit and difficulties in slaking.
Workmen should wear protective clothing and goggles to protect the skin and eyes, as
lime dust and hot slurry can cause severe burns. Areas contacted by lime should be
washed immediately. Lime should not be mixed with chemicals which have water of
hydration. The lime will be slaked by the water of hydration causing excessive temperature
rise and possibly explosive conditions. Conveyors and bins used for more than one chemical
should be thoroughly cleaned before switching chemicals.
10.3.1.2 General Design Considerations
Pebble quicklime, all passing a 3/4-in screen and not more than 5% passing a No. 100
screen, is normally specified because of easier handling and less dust. Hopper agitation
10-27
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is generally not required with the pebble form. Published slaker capacity ratings require
"soft or normally burned" limes which provide fast slaking and temperature rise, but poorer
grades of limes may also be satisfactorily slaked by selection of the appropriate slaker
retention time and capacity.
10.3.1.3 Storage
Storage of bagged lime should be in a dry place, and preferably elevated on pallets to
avoid absorption of moisture. System capacities often make the use of bagged quicklime
impractical. Maximum storage period is about 60 days.
Bulk lime is stored in air-tight concrete or steel bins having a 60° slope on the bin outlet.
Bulk lime can be conveyed by conventional bucket elevators and screw, belt, apron,
drag-chain, and bulk conveyors of mild steel construction. Pneumatic conveyors subject
the lime to air-slaking and particle sizes may be reduced by attrition. Dust collectors
should be provided on manually and pneumatically filled bins.
10.3.1.4 Feeding Equipment
A typical lime storage and feed system is illustrated in Figure 10-7. Quicklime feeders
are usually limited to the belt or loss-in-weight gravimetric types because of the wide
variation of the bulk density. Feed equipment should have an adjustable feed range of
at least 20:1 to match the operating range of the associated slaker. The feeders should
have an over-under feed rate alarm to immediately warn of operation beyond set limits
of control. The feeder drive should be instrumented to be interrupted in the event of
excessive temperature in the slaker compartment.
Lime slakers for wastewater treatment should be of the continuous type, and the major
components should include one or more slaking compartments, a dilution compartment,
a grit separation compartment and a continuous grit remover. Commercial designs vary
in regard to the combination of water to lime, slaking temperature, and slaking time,
in obtaining the "milk of lime" suspensions.
The "paste" type slaker admits water as required to maintain a desired mixing viscosity.
This viscosity therefore sets the operating retention time of the slaker. The paste slaker
usually operates with a low water to lime ratio (approximately 2:1 by weight), elevated
temperature, and five minute slaking time at maximum capacity.
The "detention" type slaker admits water to maintain a desired ratio with the lime, and
therefore the lime feed rate sets the retention time of the slaker. The detention slaker
operates with a wide range of water to lime ratios (2.5:1 and 6:1), moderate temperature,
and a 10-minute slaking time at maximum capacity. A water to lime ratio of from 3.5:1
to 4:1 is most often used. The operating temperature in lime slakers is a function of
the water to lime ratio, lime quality, heat transfer, and water temperature. Lime slaking
10-28
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FIGURE 10-7
TYPICAL LIME FEED SYSTEM
NOTE: VAPOR REMOVER
NOT SHOWN FOR CLARITY
-DUST C01LECTOR
FILL PIPE (PNEUMATIC)
BULK STORAGE
BIN
BIN GATE
FLEX IBLE
CONNECTION
SCALE
FEEDER
SOLENOID
VALVE,
OR SAMPLE CHUTE
ROTAMETERS
FLOW RECORDER
WITH PACING
TRANSMITTER,
RO TOD IP-TYPE
FEEDER
pH RECORDER
CONTROLLER
GRAVITY FEED
RECIRCULATION
BACK
PRESSURE
VALVE
10-29
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evolves heat in hydrating the CaO to Ca(OH)2 and, therefore, vapor removers are required
for feeder protection.
10.3.1.5 Piping and Accessories
Lime slurry should be transported by gravity in open channels wherever possible. Piping
channels and accessories may be rubber, iron, steel, concrete, and plastics. Glass tubing,
such as that in rotameters, will cloud rapidly and therefore should not be used. Any
abrupt directional changes in piping should include plugged tees or crosses to allow
rodding-out of deposits. Long sweep elbows should be provided to allow the piping to
be cleaned by the use of a cleaning "pig". Daily cleaning is desirable.
Milk of lime transfer pumps should be of the open impeller centrifugal type. Pumps
having an iron body and impeller with bronze trim are suitable for this purpose. Rubber
lined pumps with rubber covered impellers are also frequently used. Make-up tanks are
usually provided ahead of centrifugal pumps to ensure a flooded suction at all times.
"Plating-out" of lime is minimized by the use of soft water in the make-up tank. Turbine
pumps and eductors should be avoided in transferring milk of lime because of scaling
problems.
10.3.1.6 Pacing and Control
Lime slaker water proportioning is integrally controlled or paced from the feeder.
Therefore, the feeder-slaker system will follow pacing controls applied to the feeder only.
As discussed previously, gravimetric feeders are adaptable to receive most standard
instrumentation pacing signals. Systems can be instrumented to allow remote pacing with
telemetering of temperature and feed rate to a central panel for control purposes.
The lime feeding system may be controlled by an instrumentation system integrating both
plant flow and pH of the wastewater after lime addition. However, it should be recognized
that pH probes require daily maintenance in this application to monitor the pH accurately.
Deposits tend to build up on the probe and necessitate frequent maintenance. The low
pH lime treatment systems (pH 9.5 to 10.0) can be more readily adapted to this method
of control than high lime treatment systems (pH 11.0 or greater) because less maintenance
of the pH equipment is required. In a closed loop pH-flow control system, milk of lime
is prepared on a batch basis and transferred to a holding tank with variable output feeders
set by the flow and pH meters to proportion the feed rate. Figure 10-7 illustrates such
a control system.
10.3.2 Hydrated Lime
10-30
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10.3.2.1 Properties and Availability
Hydrated lime, Ca(OH)2, has a working density of 20 to 50 Ib/ft3, tends to flood the
feeder, and will arch in storage bins if packed. The molecular weight is 74.08. The
dust and slurry of hydrated lime are caustic in nature. The 1973 cost of bulk hydrated
lime varies from $20 to $22 per ton. Bagged lime is available but increases the cost
about $5 per ton. The availability of hydrated lime may be determined by contracting
manufacturers listed in a subsequent section. The pH of a saturated hydrated lime solution
is the same as that given for quicklime.
10.3.2.2 General Design Considerations
Hydrated lime is slaked lime and needs only enough water added to form milk of lime.
Wetting or dissolving chambers are usually designed to provide 5 minutes detention with
a ratio of 0.5 Ib per gallon of water or 6% slurry at the maximum feed rate. Hydrated
lime is usually used where maximum feed rates do not exceed 250 Ib/hour. Hydrated
lime and milk of lime will irritate the eyes, nose, and respiratory system and will dry
the skin. Affected areas should be washed with water.
10.3.2.3 Storage
Information given for quicklime also applies to hydrated lime except that bin agitation
must be provided. Bulk bin outlets should be provided with non-flooding rotary feeders
and with hopper slopes of 65%.
10.3.2.4 Feeding Equipment
Volumetric or gravimetric feeders may be used, but volumetric feeders are usually selected
only for installations where comparatively low feed rates are required. Dilution does not
appear to be important; therefore, control of the amount of water used in the feeding
operation is not considered necessary. Inexpensive hydraulic jet agitation may be furnished
in the wetting chamber of the feeder as an alternative to mechanical agitation. The jets
should be sized for the available water supply pressure to obtain proper mixing.
10.3.2.5 Piping and Accessories
Piping and accessories as described for quicklime are also appropriate for hydrated lime.
10.3.2.6 Pacing and Control
Controls as listed for dry alum apply to hydrated lime. Hydraulic jets should operate
continuously and only shut off when the feeder is taken out of service. Control of the
feed rate with pH as well as pacing with the plant flow may be used with hydrated lime
as well as quicklime.
10-31
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10.3.3 Estimated Initial Costs of Adding Lime Feed Facilities to Existing
Wastewater Treatment Plants
Cost estimates of chemical storage and feeding equipment have been prepared based on
the use of hydrated lime for a 1 mgd plant, and pebble quicklime for 10 and 100 mgd
plants. Lime feed rates are based on a dosage of 150 mg/1 and allow for peak rates
of twice this capacity. Storage was provided for at least 15 days at the average rate
although it may be desirable to provide 30 days for dry chemicals. This storage time
is arbitrary and will vary at each installation depending on the distance to and the reliability
of the source of chemical supply. Piping and buildings to house the feeding equipment
are not included in the estimates. The estimated costs of steel bins with dust collector
vents and filling accessories are included for all three plants. The chemical feeding
equipment can be installed in existing buildings at many locations. In all estimates, the
costs include an allowance for the contractor's installation, overhead, and profit, and an
allowance of 20% of the construction cost of engineering and contingencies. The estimated
costs for each size plant are shown in Table 10-8. These figures correspond to an EPA
Construction Cost Index of 185.
TABLE 10-8
COST OF LIME FEED FACILITIES
Additional Cost of
Estimated Flow and pH
Plant Size Base Bid Controls
(mgd) ($) ($)
1 19,000 5,000 24,000
10 95,000 15,000 110,000
100 365,000 100,000 475,000
Equipment for the 1.0 mgd plant includes two volumetric feeder systems with manually
loaded bins, dissolving chambers and accessories. Equipment for the 10.0 mgd plant
includes two gravimetric feeder-slaker systems. Steel bins are estimated to hold 1-1/2
truckloads of quicklime each, and it is assumed that the truck will be equipped with
a pneumatic blower. Equipment for the 100.00 mgd plant includes four gravimetric
feeder-slaker systems. The base estimate includes the bin gate and feeder-slaker accessories
and controls.
10-32
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Costs for the flow and pH controls include holding tanks with mixers, and slurry feeders
with controls. Flow and pH metering instruments are not included.
10.4 Compounds for pH Adjustment
10.4.1 Soda Ash
10.4.1.1 Properties and Availability
Soda ash, Na2CO^, is available in two forms. Light soda ash has a bulk density range
of 35 to 50 lb/ft^ and a working density of 41 lb/ft3. Dense soda ash has a density
range of 60 to 76 lb/ft3 and a working density of 63 lb/ft3. The pH of a 1% solution
of soda ash is 1 1.2.
The molecular weight of soda ash is 106. Commercial purity ranges from 98 to 99%.
The viscosities of sodium carbonate solutions are given in Figure 10-8. Soda ash by itself
is not particularly corrosive, but in the presence of lime and water, caustic soda is formed
which is quite corrosive.
Soda ash is available in bulk by truck, box car and hopper car, and in 100 Ib bags from
locations listed in a subsequent section. The 1974 price for soda ash ranges from $50
to $60 per ton, F.O.B. the point of manufacture; however, prices vary substantially between
manufacturers and should be obtained from the closest manufacturers or local distributors.
10.4.1.2 General Design Considerations
Dense soda ash is generally used in municipal applications because of superior handling
characteristics. It has little dust, good flow characteristics, and will not arch in the bin
or flood the feeder. It is relatively hard to dissolve and ample dissolver capacity must
be provided. Normal practice calls for 0.5 Ib of dense soda ash per gallon of water or
a 6% solution retained for 20 minutes in the dissolver.
The dust and solution are irritating to the eyes, nose, lungs and skin and therefore general
precautions should be observed and the affected areas should be washed promptly with
water.
10.4.1.3 Storage
Soda ash is usually stored in steel bins and where pneumatic filling equipment is used,
bins should be provided with dust collectors. Bulk and bagged soda ash tend to absorb
atmospheric carbon dioxide and water forming the less active sodium bicarbonate
(NaHCOj). Material recommended for unloading facilities is steel.
10-33
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eo
UJ
CO
CO
o
o
en
FIGURE 10-8
VISCOSITY OF SODA ASH SOLUTIONS
(COURTESY PPG INDUSTRIES INC., CHEMICAL DIV.)
6.0 —
5.0 -
10-34
-------
10.4.1.4 Feeding Equipment
Feed equipment as described for dry alum is suitable for soda ash. Dissolving of soda
ash may be hastened by the use of warm dissolving water. Mechanical or hydraulic jet
mixing should be provided in the dissolver.
10.4.1.5 Piping and Accessories
Materials of construction for piping and accessories should be iron, steel, rubber, and
plastics.
10.4.1.6 Pacing and Control
Controls as discussed for dry alum apply also to soda ash equipment.
10.4.2 Liquid Caustic Soda
Anhydrous caustic soda (NaOH) is available but its use is generally not considered practical
in water and wastewater treatment applications. Consequently, only liquid caustic soda
is discussed below.
10.4.2.1 Properties and Availability
Liquid caustic soda is shipped at two concentrations, 50% and 73% NaOH. The densities
of the solutions as shipped are 12.76 Ib/gal for the 50% solution and 14.18 Ib/ga for
the 73% solution. These solutions contain 6.38 Ib/gal NaOH and 10.34 Ib/gal NaOH,
respectively. The crystallization temperature is 53°F for the 50% solution and 165°F
for the 73% solution. The molecular weight of NaOH is 40. Viscosities of various caustic
soda solutions are presented in Figure 10-9. The pH of a 1% solution of caustic soda
is 12.9.
Truckload lots of 1,000 to 4,000 gallons are available in the 50% concentration only.
Both shipping concentrations can be obtained in 8,000, 10,000, and 16,000 gallon carload
lots. Tank cars can be unloaded through the dome eduction pipe using air pressure or
through the bottom valve by gravity or by using air pressure or a pump. Trucks are
usually unloaded by gravity or with air pressure or a truck mounted pump.
Major producers of caustic soda and their respective plant locations are listed in a
subsequent section. The current price for liquid caustic soda ranges from about $75 to
$80 per ton (NaOH), F.O.B. the point of manufacture.
10-35
-------
co
o
o
200
100
80
60
40
20
10
8
6
4
1
0.8
0.6
0-4
0.2
FIGURE 10-9
VISCOSITY OF CAUSTIC SODA SOLUTIONS
(COURTESY OF HOOKER CHEMICAL CO.)
I I
I I
I
I
I
I
I
I
I
I
60 70 80 90 100 110 120 130 140 150 160 170 180 190 200
TEMPERATURE, °F
10-36
-------
10.4.2.2 General Design Considerations
Liquid caustic soda is received in bulk shipments, transferred to storage and diluted as
necessary for feeding to the points of application. Caustic soda is poisonous and is
dangerous to handle. U.S. Department of Transportation Regulations for "White Label"
materials must be observed. However, if handled properly caustic soda poses no particular
industrial hazard. To avoid accidental spills, all pumps, valves, and lines should be checked
regularly for leaks. Workmen should be thoroughly instructed in the precautions related
to the handling of caustic soda. The eyes should be protected by goggles at all times
when exposure to mist or splashing is possible. Other parts of the body should be protected
as necessary to prevent alkali burns. Areas exposed to caustic soda should be washed
with copious amounts of water for 15 minutes to 2 hours. A physician should be called
when exposure is severe. Caustic soda taken internally should be diluted with water or
milk and then neutralized with dilute vinegar or fruit juice. Vomiting may occur
spontaneously but should not be induced except on the advice of a physician.
10.4.2.3 Storage
Liquid caustic soda may be stored at the 50% concentration. However, at this solution
strength, it crystallizes at 53°F. Therefore, storage tanks must be located indoors or
provided with heating and suitable insulation if outdoors. Because of its relatively high
crystallization temperature, liquid caustic soda is often diluted to a concentration of about
20% NaOH for storage. A 20% solution of NaOH has a crystallization temperature of
about -20°F. Recommendations for dilution of both 73% and 50% solutions should be
obtained from the manufacturer, because special considerations are necessary.
Storage tanks for liquid caustic soda should be provided with an air vent for gravity flow.
The storage capacity should be equal to 1-1/2 times the largest expected delivery, with
an allowance for dilution water, if used, or 2 weeks supply at the anticipated feed rate,
whichever is greater. Tanks for storing 50% solution at a temperature between 75°F
and 140°F may be constructed of mild steel. Storage temperatures above 140°F require
more elaborate materials selection and are not recommended. Caustic soda will tend to
pick up iron when stored in steel vessels for extended periods. Subject to temperature
and solution strength limitations, rubber, 316 stainless steel, nickel, nickel alloys, or plastics
may be used when iron contamination must be avoided.
10.4.2.4 Feeding Equipment
Further dilution of liquid caustic soda below the storage strength may be desirable for
feeding by volumetric feeders. Feeding systems as described for liquid alum generally
apply to caustic soda with appropriate selection of materials of construction. A typical
system schematic is shown in Figure 10-10. Feeders will usually include materials such
as ductile iron, stainless steels, rubber, and plastics.
10-37
-------
FIGURE 10-10
TYPICAL CAUSTIC SODA FEED SYSTEM
-TRUCK FILL LINE
DILUTION
WATER
SODIUM HYDROXIDE
STORAGE TANK
VENT, OVERFLOW
AND DRAIN
VENT, OVERFLOW
AND DRAIN
MIXER
SAMPLE TAP
ODIUM HYDROXIDE
FEEDER
POINT OF
APPLICATION
10-38
-------
10.4.2.5
Piping and Accessories
Transfer lines from the shipping unit to the storage tank should be spiral-wire-bound
neoprene or rubber hose, solid steel pipe with swivel joints, or steel hose. Because caustic
soda attacks glass, use of glass materials should be avoided. Other miscellaneous materials
for use with liquid caustic soda feeding and handling equipment are listed below (7) in
Table 10-9:
TABLE 10-9
MATERIALS SUITABLE FOR CAUSTIC SODA SERVICE
Components
Rigid Pipe
Flexible Connections
Diluting Tees
Fittings
Permanent Joints
Unions
Valves - Non-leaking (Plug)
Body
Plug
Pumps (Centrifugal)
Body
Impeller
Packing
Storage Tanks
Recommended Materials
for Use With 50% NaOH
Up to 140°F
Standard Weight Black Iron
Rigid Pipe with Ells or Swing
Joints, Stainless Steel or
Rubber Hose
Type 304 Stainless Steel
Steel
Welded or Screwed Fittings
Screwed Steel
Steel
Type 304 Stainless Steel
Steel
Ni-Resist
Blue Asbestos
Steel
10.4.2.6 Pacing and Control
Controls as listed for liquid alum also apply to liquid caustic soda equipment.
10-39
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10.5 Carbon Dioxide
10.5.1 Properties and Availability
Carbon dioxide, CC^, is available for use in wastewater treatment plants in gas and liquid
form. The molecular weight of carbon dioxide is 44. Dry carbon dioxide is not chemically
active at normal temperatures and is a non-toxic safe chemical; however, the gas displaces
oxygen and adequate ventilation of closed areas should be provided. Solutions of
carbon dioxide in water are very reactive chemically and form carbonic acid. Saturated
solutions of carbon dioxide have a pH of 4.0 at 68°F.
The gas form may be produced on the treatment plant site by scrubbing and compressing
the combustion product of lime recalcining furnaces, sludge furnaces, or generators used
principally for the production of carbon dioxide gas only. These generators are usually
fired with combustible gases, fuel oil, or coke and have carbon dioxide yields as shown
in Table 10-10 (8).
TABLE 10-10
CARBON DIOXIDE YIELDS OF COMMON FUELS
Fuel Quantity Yield
Ib
Natural Gas 1,000 ft3 115
Coke 1 Ib 3
Kerosene 1 gal 20
Fuel Oil (No. 2) 1 gal 23
Propane 1,000 ft3 352
Butane 1,000 ft3 454
The gas forms, as generated at the plant site, usually have a carbon dioxide content of
between 6% and 18% depending on the source and efficiency of the producing system.
The liquid form is available from commercial suppliers in 20 to 50 Ib cylinders, 10 to
20 ton trucks and 30 to 50 ton rail cars. The commercial liquid form has a minimum
carbon dioxide content of 99.5%.
10-40
-------
1974 prices range from S33-S36 per ton for 3,000 tons per year and over, to $71-$74
per ton for a quantity of 150 tons per year. These prices include an allowance for freight
within a 100 mile radius of the point of manufacture. Another $6 or $7 per ton may
be added for each additional 100 miles to the point of destination. Major producers
of commercial carbon dioxide are listed in a subsequent section.
10.5.2 General Design Considerations
Recovery of carbon dioxide from recalcining furnaces or incinerators is the least expensive
source, but maintenance of stack gas systems is likely to be extensive because of the
corrosive nature of the wet gas and the presence of particulate matter. Scrubber systems
are required to clean the stack gas and especially designed gas compressors are necessary
to provide the process injection pressure.
Pressure generators and submerged burners require less maintenance because the system
pressure is established by compressors or blowers handling dry air or gas. On-site generating
units have a limited range of carbon dioxide production as compared with the liquid storage
and feed system, and therefore may require multiple units.
The liquid carbon dioxide storage and feed system generally includes a temperature-pressure
controlled, bulk storage tank, an evaporation unit, and a gas feeder to meter the gas.
Solution feeders, similar in construction to chlorinators, may also be used to feed carbon
dioxide.
10.5.3 Storage
This section applies only to use of commercial liquid carbon dioxide. Liquid system
capacities encountered in wastewater treatment usually usually require on-site bulk storage
units. Standard pre-packaged units are available, ranging in size from 3/8 to 50 tons
capacity, and are furnished with temperature-pressure controls to maintain approximately
300 lb/in2 at 0°F conditions. The typical package unit con tains refrigeration, vaporization,
safety and control equipment. The units are well insulated and protected for outdoor
location. The gas from the evaporation unit usually passes through two stages of pressure
reduction before entering the gas feeder to prevent the formation of dry ice.
10.5.4 Feeding Equipment
Feeding systems for the stack gas source of carbon dioxide consist of simple valving
arrangements, for admitting varying quantities of make-up air to the suction side of the
constant volume compressors, or for venting excess gas on the compressor discharge. A
typical system is described by Gulp and Gulp (9).
10-41
-------
Pressure generators and submerged burners are regulated by valving arrangements on the
fuel and air supply. Generation of carbon dioxide by combustion is usually difficult
to control, requires frequent operator attention and demands considerable maintenance
over the life of the equipment, when compared to liquid CC>2 systems.
Commercial liquid carbon dioxide is becoming more generally used because of its high
purity, the simplicity and range of feeding equipment, ease of control, and smaller, less
expensive piping systems. After vaporization, carbon dioxide with suitable metering and
pressure reduction may be fed directly to the point of application as a gas. Metering
of direct fed pressurized gas is difficult due to high adiabatic expansion characteristics
of the gas. Also, direct feed requires extremely fine bubbles to insure the gas goes into
solution; this, in turn, can lead to scaling problems. However, vacuum operated, solution
type gas feeders are preferred. Such feeders generally include safety devices and operating
controls in a compact panel housing, with materials of construction suitable for carbon
dioxide service. Absorption of carbon dioxide in the injector water supply approaches
100% when a ratio of 1.0 Ib of gas to 60 gal of water is maintained.
10.5.5 Piping and Accessories
Mild steel piping and accessories are suitable for use with cool, dry, carbon dioxide. Hot,
moist gases, however, require the use of type 316 stainless steel or plastic materials. Plastics
or FRP pipe are generally used for solution piping and diffusers. Diffusers should be
submerged at least 8 ft, and preferably deeper, to assure complete absorption of the gas.
10.5.6 Pacing and Control
Standard instrument signals and control components can be used to pace or control carbon
dioxide feed systems.
Using stack gas as the source of carbon dioxide, the feed rate can be controlled by proper
selection and operation of compressors, by manual control of vent or bleed valves, or
by automatic control of such valves by a pH meter-controller system.
In commercial carbon dioxide feed systems, solution feeders may function as controllers
and can be paced by instrument signals from pH monitors and plant flow meters.
In feeding commercial carbon dioxide directly to the point of application as a gas, a
differential pressure transmitter and a control valve may function as the primary elements
of a control system. Standard instrument signals may be used to pace or control the
rate of carbon dioxide feed.
Carbon dioxide generators are difficult to pace or control other than by manual or
automatic operation of vent or bleed valves that waste a portion of the produced gas
according to the plant requirements.
10-42
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10.6 Polymers
10.6.1 Dry Polymers
10.6.1.1 Properties and Availability
Types of polymers vary widely in characteristics. Manufacturers should be consulted for
properties, availability, and cost of the polymer being considered. References are available
that indicate the types and characteristics of polymers available (10, 11). Bulk shipments
are generally not desirable. Polymers are available in a variety of container or package
sizes.
10.6.1.2 General Design Considerations
Dry polymer and water must be blended and mixed to obtain a recommended solution
for efficient action. Solution concentrations vary from fractions of a percent up.
Preparation of the stock solution involves wetting of the dry material and usually an aging
period prior to application. Solutions can be very viscous, and close attention should
be paid to piping size and length and pump selections. Metered solution is usually diluted
just prior to injection to the process to obtain better dispersion at the point of application.
10.6.1.3 Storage
General practice for storage of bagged dry chemicals should be observed. The bags should
be stored in a dry, cool, low humidity area and used in proper rotation, i.e., first in,
first out.
Solutions are generally stored in type 316 stainless steel, FRP, or plastic lined tanks.
10.6.1.4 Feeding Equipment
Two types of systems are frequently combined to feed polymers. The solution preparation
system includes a manual or automatic blending system with the polymer dispensed by
hand or by a dry feeder to a wetting jet and then to a mixing-aging tank at a controlled
ratio. The aged polymer is transported to a holding tank where metering pumps or rotodip
feeders dispense the polymer to the process. A schematic of such a system is shown
by Figure 10-11. It is generally advisable to keep the holding or storage time of polymer
solutions to a minimum, 1 to 3 days or less, to prevent deterioration of the product.
10.6.1.5 Piping and Accessories
Selection must be made after determination of the polymer; however, type 316 stainless
steel or plastics are generally used.
10-43
-------
FIGURE 10-11
MANUAL DRY POLYMER FEED SYSTEM
WATER SUPPLY
— DRY
FEEDER
DISPERSER
MIXER
DISSOLVING-AGING
TANK
HOLDING TANK
SOLUTION FEEDER
POINT OF
APPLICATION
10-44
-------
10.6.1.6 Pacing and Control
Controls as listed for liquid alum apply to the control of liquid dispersing feeders.
The solution preparation system may be an automatic batching system, as shown by the
schematic on Figure 10-12, that fills the holding tank with aged polymer as required by
level probes. Such a system is usually provided only at large plants. Unitized solution
preparation units are available, but have a limited capacity.
10.6.2 Liquid Polymers
10.6.2.1 Properties and Availability
As with dry polymers, there is a wide variety of products, and manufacturers should be
consulted for specific information.
10.6.2.2 General Design Considerations
Liquid systems differ from the dry systems only hi the equipment to blend the polymer
with water to prepare the solution. Liquid solution preparation is usually a hand-batching
operation with manual filling of a mixing-aging tank with water and polymer.
10.6.2.3 Feeding Equipment
Liquid polymers need no aging and simple dilution is the only requirement for feeding.
The dosage of liquid polymers may be accurately controlled by metering pumps or rotodip
feeders.
The balance of the process is generally the same as described for dry polymers.
10.6.3 Estimated Initial Costs of Adding Polymer Feed Facilities to
Existing Wastewater Treatment Plants
Cost estimates of solution preparation and feeder equipment have been prepared, based
on the use of dry polymer, for 1, 10, and 100 mgd plants. Chemical feed equipment
was sized to have a capacity sufficient to prepare and feed a 0.25% stock solution at
a dosage of 1 mg/1. Piping and buildings to house the feeding equipment and store the
bags are not included. At many locations, the chemical feeding equipment can be installed
in existing buildings. All estimated costs include an allowance for the contractor's
installation, overhead, and profit plus an allowance of 20% of the construction cost for
engineering and contingencies. The estimated costs for each size plant are shown in
Table 10-11. These figures correspond to an EPA Construction Cost Index of 185:
10-45
-------
FIGURE 10-12
AUTOMATIC DRY POLYMER FEED SYSTEM
HOT
MATER
SOLENOID /SCALE
/ VALVE
DISPERSES
NOTE: CONTROL t INSTRUMENTATION
WIRIN8 IS NOT SHOW
POINT OF APPLICATION
SOLUTI ON
FEEDERS.^
^"~-—
••
\
^••i^
n
HOLDING TANK
1—
TRANSFER PUMP
LEVEL PROBE
10-46
-------
TABLE 10-11
COST OF POLYMER FEED FACILITIES
Plant Size Estimated Cost
(mgd) ($)
1 7,000
10 31,000
100 225,000
The system for the 1 mgd plant was based on the polymer being manually fed to the
mixing tank, with mixing and solution feed equipment arranged for manual control. Two
independent systems of tanks and feeders are included. The system for the 10 mgd plant
includes two volumetric dry feeders discharging to two mixing tanks, arranged for batch
control. The mixing tanks discharge to a single holding tank from which two solution
feeders take their supply for application to the process. The system for the 100 mgd
plant includes four volumetric dry feeders complete with steel day hoppers, dust collectors,
bin gates and flexible connectors. The system operation is automatic for the preparation
and transfer of aged polymer solution from four mixing tanks to two holding tanks. Ten
solution feeders meter the polymer to the treatment process.
10.7 Chemical Suppliers
10.7.1 Metal Salts and Lime
The first and largest list of suppliers concerns metal salts and lime. It is divided
geographically into states, which are tabulated alphabetically. Then, under each state,
production locations are listed for the following alphabetized chemicals:
alum (dry and/or liquid)
ferric chloride
ferric sulfate
ferrous sulfate
lime (high calcium and/or dolomitic)
sodium aluminate (dry and/or liquid)
10-47
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Table 10-12 summarizes extent and location of presently available sources of metal salts
and lime:
TABLE 10-12
GEOGRAPHICAL SUMMARY OF SUPPLIERS OF METAL SALTS AND LIME
Location Chemical
25 States Alum
Michigan Ferric chloride
Tennessee Ferric sulfate
Ga., Md., Mo., NJ. Ferrous sulfate
32 states Lime
Ark., 111., La., Ga., Mo. Sodium aluminate
Table 10-13 offers a detailed listing of metal salt and lime suppliers. Persons interested
in contacting chemical suppliers may direct inquiries to either the point of production
or to a nearby sales office for the chemical company involved. This list does not include
potential or established suppliers of waste pickle liquor. Nor are pH adjusting chemicals
included here; these are normally available through most major chemical supply firms.
10-48
-------
TABLE 10-13
SUPPLIERS OF METAL SALTS AND LIME
Location
Manufacturer
Chemical
ALABAMA
Coosa Pines
Demopolis
Mobile
Naheola
Allgood
Keystone
Landmark
Monte vallo
Roberta
Saginaw
Siluria
American Cyanamid
American Cyanamid
American Cyanamid
Stauffer
Cheney Lime & Cement Co.
Martin Marietta Cement,
Southern Division
Cheney Lime & Cement Co.
U.S. Gypsum Co.
Southern Cement Co.,
Div. Martin Marietta
Corp.
Longview Lime Co., Div.
Woodward Co., Div.
Mead Corp.
Alabaster Lime Co.
Alum, Liquid
Alum, Liquid
Alum, Liquid and Dry
Alum, Liquid
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
ARIZONA
Douglas
Globe
Nelson
Paul Lime Plant, Inc.
Hoopes & Co.
U.S. Lime Div., The
Flintkote Co.
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
ARKANSAS
Pine Bluff
Batesville
Bauxite
Allied
Batesville White Lime
Co., Div. Rangaire
Corp.
Reynolds Chemicals
Alum, Liquid
Lime, High Calcium
Sodium Aluminate,
Dry
10-49
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SUPPLIERS OF METAL SALTS AND LIME
Location
Manufacturer
Chemical
CALIFORNIA
Bay Point
Antioch
El Segundo
Richmond
Vernon
City of Industry
Diamond Springs
Lucerne Valley
Richmond
Salinas
Westend
Allied
Imperial West Chemicals
Allied
Stauffer
Stauffer
U.S. Lime Div., The
Flintkote Co.
Diamond Springs Lime Co.
Pfizer, Inc., Minerals,
Pigments & Metals
Div.
U.S. Lime Div., The
Flintkote Co.
Kaiser Aluminum &
Chemical Corp.
Stauffer Chemical Co.
Alum, Liquid and Dry
Ferric Chloride
Alum, Liquid
Alum, Liquid
Alum, Liquid
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Lime, Dolomitic
Lime, High Calcium
COLORDAO
Denver
Ft. Morgan
Allied
Great Western Sugar Co.
Alum, Liquid and Dry
Lime, High Calcium
CONNECTICUT
Canaan
Pfizer, Inc., Minerals
Pigments & Metals
Div.
Lime, Dolomitic
FLORIDA
Fernandina Beach
Jacksonville
Port St. Joe
Tennessee Corp.
Allied
Allied
Alum, Liquid
Alum, Liquid
Alum, Liquid
10-50
-------
SUPPLIERS OF METAL SALTS AND LIME
Location
FLORIDA (cont.)
Brooksville
Sumterville
Manufacturer
Chemical Lime Co.
Dixie Lime and Stone Co.
Chemical
Lime, High Calcium
Lime, High Calcium
GEORGIA
Atlanta (2)
Augusta
Cedar Springs
Macon
Savannah
Savannah
Vinings
Burris; Allied
Tennessee Corp.
Tennessee Corp.
Allied
Allied
American Cyanamid Co.
Vinings Chemical Co.
Alum, Liquid and Dry
Alum, Liquid and Dry
Alum, Liquid
Alum, Liquid
Alum, Liquid
Ferrous Sulfate
Sodium Aluminate,
Liquid
ILLINOIS
E. St. Louis
Joliet
Marblehead
McCook
Quincy
So. Chicago
Thornton
Chicago
Allied
American Cyanamid
Marblehead Lime Co.
Vulcan Materials Co.
Marblehead Lime Co.
Marblehead Lime Co.
Marblehead Lime Co.
Nalco Chemical Company
Alum, Liquid and Dry
Alum, Liquid and Dry
Lime, High Calcium
Lime, Dolomitic
Lime, High Calcium
Lime, High Calcium
Lime, Dolomitic
Sodium Aluminum,
Liquid and Dry
INDIANA
Buffington
Marblehead Lime Co.
Lime, High Calcium
10-51
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SUPPLIERS OF METAL SALTS AND LIME
Location
IOWA
Davenport
Manufacturer
Linwood Stone Products
Co., Inc.
Chemical
Lime, High Calcium
KENTUCKY
Carntown
Black River Mining Co.
Lime, High Calcium
LOUISIANA
Bastrop
Baton Rouge
Monroe
New Orleans
Springhill
Morgan City
New Orleans
Stauffer
Stauffer
Allied
Allied
Stauffer
Pelican State Lime Corp.
U.S. Gypsum Co.
Alum, Liquid
Alum, Liquid
Alum, Liquid
Alum, Liquid and Dry
Alum, Liquid
Lime, High Calcium
Lime, High Calcium
MAINE
Searsport
Northern
Alum, Liquid and Dry
MARYLAND
Baltimore
Baltimore (3)
Woodsboro
Olin
Byproducts Processing
Co., Inc.;
Glidden Co.;
Cosmin Corp.;
S.W. Barrick & Sons, Inc.
Alum, Dry
Ferrous Sulfate
Ferrous Sulfate
Ferrous Sulfate
Lime, High Calcium
10-52
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SUPPLIERS OF METAL SALTS AND LIME
Location
MASSACHUSETTS
Adams
Salem
Adams
Lee
Manufacturer
Holland
Hamblet & Hayes
Pfizer, Inc., Minerals
Pigments & Metals
Div.
Lee Lime Corp.
Chemical
Alum, Liquid
Alum, Liquid and Dry
Lime, High Calcium
Lime, Dolomitic
MICHIGAN
Detroit
Escanaba
Kalamazoo (2)
Midland
Wyandotte
Detroit
Ludington
River Rouge
Cloquet
Pine Bend
Duluth
Allied
American Cyanamid
Allied; American Cyanamid
Dow Chemical Co.
Pennwalt Corp.
Detroit Lime Co.
Dow Chemical Co.
Marblehead Lime Co.
American Cyanamid
North Star
Cutler Magner Co.
Alum, Liquid
Alum, Liquid
Alum, Liquid
Ferric Chloride
Ferric Chloride
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Alum, Liquid
Alum, Liquid and Dry
Lime, High Calcium
MISSISSIPPI
Monticello
Vicksburg
American Cyanamid
Allied
Alum, Liquid
Alum, Liquid
10-53
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SUPPLIERS OF METAL SALTS AND LIME
Location
Manufacturer
Chemical
MISSOURI
St. Louis
Bonne Terre
Hannibal
Ste. Genevieve
Springfield
Kansas City
National Lead
Valley Mineral Products
Co.
Marblehead Lime Co.
Mississippi Lime Co.
Ash Grove Cement Co.
Conservation Chemical Co.
Ferrous Sulfate
Lime, Dolomitic
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Sodium Aluminate,
Liquid
NEVADA
Apex
Henderson
McGill
Sloan
U.S. Lime Div., The
Flintkote Co.
U.S. Lime Div., The
Flintkote Co.
Morrison-Weatherly Corp.
U.S. Lime Div., The
Flintkote Co.
Lime, High Calcium
Lime, Dolomitic and
High Calcium
Lime, High Calcium
Lime, Dolomitic and
High Calcium
NEW JERSEY
Newark
Warners
Sayreville
Newton
Essex
American Cyanamid
National Lead
Limestone Products
Corp. of America
Alum, Liquid and Dry
Alum, Liquid and Dry
Ferrous Sulfate
Lime, High Calcium,
Hydrate Only
10-54
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SUPPLIERS OF METAL SALTS AND LIME
Location
NORTH CAROLINA
Acme
Plymouth
Manufacturer
Wright
American Cyanamid
Chemical
Alum, Liquid
Alum, Liquid
OHIO
Chillicothe
Cleveland
Hamilton
Middletown
Ashtabula
Carey
Cleveland
Delaware
Genoa
Gibson burg (2)
Huron
Marble Cliff
Millersville
Woodville (2)
Allied
Allied
American Cyanamid
Allied
Union Carbide Olefins Co.
National Lime & Stone Co.
Cuyahoga Lime Co.
MCQ Co.
U.S. Gypsum Co.
Pfizer, Inc., Minerals,
Pigments & Metal
Div.; National
Gypsum Co.
Huron Lime Co.
MCQ Co.
I.E. Baker Co.
Ohio Lime Co.; Standard
Lime & Refractories
Div., Martin
Marietta Corp.
Alum, Liquid
Alum, Liquid and Dry
Alum, Liquid and Dry
Alum, Liquid
Lime, High Calcium
Lime, Dolomitic
Lime, High Calcium
Lime, High Calcium
Lime, Dolomitic
Lime, Dolomitic
Lime, High Calcium
Lime, High Calcium
Lime, Dolomitic
Lime, Dolomitic
OKLAHOMA
Marble City
Sallisaw
St. Clair Lime Co.
St. Clair Lime Co.
Lime, High Calcium
Lime, High Calcium
10-55
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SUPPLIERS OF METAL SALTS AND LIME
Location
OREGON
North Portland
Baker
Portland
Manufacturer
Stauffer
Chemical Lime Co. of
Oregon
Ash Grove Cement Co.
Chemical
Alum, Liquid and Dry
Lime, High Calcium
Lime, High Calcium
PENNSYLVANIA
Johnsonburg
Marcus Hook
Newell
Annville
Bellefonte (2)
Branchton
Devault
Pleasant Gap
Plymouth Meeting
Allied
Allied
Allied
Bethlehem Mines Corp.
National Gypsum Co.;
Warner Co.
Mercer Lime & Stone Co.
Warner Co.
Chemical Lime Inc.
G.&W.H. Corson, Inc.
Alum, Liquid
Alum, Liquid and Dry
Alum, Liquid
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Lime, Dolomitic
Lime, High Calcium
Lime, Dolomitic
SOUTH CAROLINA
Catawba
Georgetown
Burris
American Cyanamid
Alum, Liquid
Alum, Liquid and Dry
SOUTH DAKOTA
Rapid City
Pete Lien & Sons, Inc.
Lime, High Calcium
10-56
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SUPPLIERS OF METAL SALTS AND LIME
Location
Manufacturer
Chemical
TENNESSEE
Chattanooga
Counce
Springfield
Copper Hill
Knoxville (2)
American Cyanamid
Stauffer
Burris
Cities Service Co.
Foote Mineral Co.;
Williams Lime
Mfg. Co.
Alum, Liquid and Dry
Alum, Liquid
Alum, Liquid
Ferric Sulfate
Lime, High Calcium
TEXAS
Houston (2)
Texas City
Blum
Cleburne
Clifton
Houston
McNeil
New Braunfels
Round Rock
San Antonio
Stauffer; Ethyl
Gulf Chemical &
Metallurgical
Round Rock Lime
Companies
Texas Lime Co., Div.
Rangaire Corp.
Chemical Lime Inc.
U.S. Gypsum Co.
Austin White Lime Co.
U.S. Gypsum Co.
Round Rock Lime
Companies
McDonough Bros., Inc.
Alum, Liquid and Dry
Ferrous Sulfate,
Ferric Sulfate
Lime, High Calcium
Lime, High Calcium
Lime, High
Lime, High
Lime, High
Lime, High
Lime, High
Calcium
Calcium
Calcium
Calcium
Calcium
Lime, High Calcium
UTAH
Grantsville
Lehi
U.S. Lime Div., The
Flintkote Co.
Rollins Mining Supplies
Co.
Lime, Dolomitic and
High Calcium
Lime, High Calcium
10-57
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SUPPLIERS OF METAL SALTS AND LIME
Location
Manufacturer
Chemical
VERMONT
Winooski
Vermont Assoc. Lime
Industries, Inc.
Lime, High Calcium
VIRGINIA
Covington
Hope well
Norfolk
Clearbrook
Kimballton (2)
Stephens City
Strasburg
Allied
Allied
Howerton Gowen
W.S. Frey Co., Inc.
Foote Mineral Co.;
National Gypsum Co.
M.J. Grove Lime Co.,
Div. The Flintkote
Co.
Chemstone Corp.
Alum, Liquid
Alum, Liquid
Alum, Liquid
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
Lime, High Calcium
WASHINGTON
Kennewick
Tacoma (2)
Vancouver
Tacoma
Allied
Allied; Stauffer
Allied
Domtar Chemicals Inc.
Alum, Liquid
Alum, Liquid
Alum, Liquid and Dry
Lime, High Calcium
WEST VIRGINIA
Riverton
Germany Valley Limestone
Div., Greer Steel Co.
Lime, High Calcium
10-58
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SUPPLIERS OF METAL SALTS AND LIME
Location Manufacturer Chemical
WISCONSIN
Manasha Allied Alum, Liquid
Wisconsin Rapids Allied Alum, Liquid
Eden Western Lime & Cement Co. Lime, Dolomitic
Green Bay Western Lime & Cement Co. Lime, High Calcium
Knowles Western Lime & Cement Co. Lime, Dolomitic
Manitowoc Rockwell Lime Co. Lime, Dolomitic
Superior Cutler-LaLiberte- Lime, High Calcium
McDougall Corp.
10.7.2 Polymer Suppliers
Table 10-14 lists names and addresses of firms offering polymer which may prove effective
in treatment approaches described herein.
TABLE 10-14
POLYMER SOURCES AND TRADE NAMES
Source Trade Name(s)
Allied Colloids, Inc. Percol
One Robinson Lane
Ridgewood, NJ 07450
Allstate Chemical Co. Allstate
Post Office Box 3040
Euclid, OH 44117
10-59
-------
POLYMER SOURCES AND TRADE NAMES
Sources
Allyn Chemical Co.
2224 Fairhill Road
Cleveland, OH 44106
American Cyanimid Co.
Berdan Avenue
Wayne, NJ 07470
Atlad Chemical Ind., Inc.
Wilmington, DE 19899
Berdell Industries
28-01 Thomson Avenue
Long Island City, NY 11101
Betz Laboratories, Inc.
Somerton Road
Trevose, PA 19047
Bond Chemicals, Inc.
1500 Brookpark Road
Cleveland, OH 44109
Brennan Chemical Co.
704 N. First Street
St. Louis, MO 63102
The Burtonite Company
Nutley, NJ 07110
Calgon Corporation
Post Office Box 1346
Pittsburgh, PA 15222
Trade Name(s)
Claron
Superfloc
Magnifloc
Sarbo
Atlasep
Berdell
Betz
Polyfloc
Bondfloc
Brenco
Burtonite
Cat-Floe
10-60
-------
POLYMER SOURCES AND TRADE NAMES
Source
Commercial Chemical
11 Patterson Avenue
Midland Park, NJ 07432
Dearborn Chemical Div.
W. R. Grace & Co.
Merchandise Mart Plaza
Chicago, IL 60654
Dow Chemical USA
Barstow Building
2020 Dow Center
Midland, MI 48640
Drew Chemical Corp.
701 Jefferson Road
Parsippany, NJ 07054
DuBois Chemicals Div.
W. R. Grace & Co.
3630 E. Kemper Road
Sharonville, OH 45241
E. I. DuPont de Nemours & Co.
Eastern Laboratory
Gibbstown, NJ 08027
Environmental Pollution Investigation
& Control, Inc.
9221 Bond Street
Overland Park, KS 66214
Trade Name(s)
Speedifloc
Aquafloc
Dowell
PEI
Purifloc
Separan XD
Drewfloc
Amerfloc
Flocculite
DuPont
Dynafloc
10-61
-------
POLYMER SOURCES AND TRADE NAMES
Source Trade Name(s)
Fabcon International Zulcar
1275 Columbus Avenue Fabcon
San Francisco, CA 94133
Henry W. Fink & Co. Kleer-Floc
6900 Silverton Avenue
Cincinnati, OH 45236
Gamlen Sybron Corp. Gamafloc
321 Victory Avenue Gamlose
S. San Francisco, CA 94080 Gamlen
Garrett-Callahan Garrett-Callahan
111 Rollins Road
Millbrae, CA
General Mills Chemicals Supercol
4620 N. 77th Street Guartec
Minneapolis, MN 55435
Hercules, Inc. Hercofloc
910 Market Street
Wilmington, DE 19899
Frank Herzl Corp. Perfectamyl
299 Madison Avenue
New York, NY 10017
ICI America, Inc. Atlasep
Wilmington, DE 19899
Illinois Water Treatment Co. Illco
840 Cedar Street
Rockford, IL 61102
10-62
-------
POLYMER SOURCES AND TRADE NAMES
Source Trade Name(s)
Kelco Company Kelgin
8225 Aero Drive Kelcosol
San Diego, CA 92123
Key Chemicals Key-Floe
4346 Tacony
Philadelphia, PA 19124
Metalene Chemical Co. Metalene
Bedford, OH 44014
The Mogul Corporation Mogul
20600 Chagrin Boulevard
Cleveland, OH 44122
Nalco Chemical Co. Nalcolyte
6216 W. 66th PI.
Chicago, IL 60638
Narvon Mining & Chemical Co. Sink-Floe
Keller Avenue & Fruitville Pike Zeta-Floc
Lancaster, PA 17604
National Starch & Chemical Corp. Floe-Aid
1700 W. Front Street Natron
Plainfield, NJ 07063
O'Brein Industries, Inc. O'B Floe
95 Dorsa Avenue
Livingston, NJ 07039
Oxford Chemical Div. Oxford-Hydro-Floc
Consolidated Foods Corp.
Post Office Box 80202
Atlanta, GA 30341
10-63
-------
POLYMER SOURCES AND TRADE NAMES
Source Trade Name(s)
Reichhold Chemicals, Inc. Aquarid
RCI Building
White Plains, NY 10602
Standard Brands Chem. Ind., Inc. Tychem
Post Office Drawer K
Dover, DE 19901
A. E. Staley Mfg. Co. Hamaco
Post Office Box 151
Decatur, IL 62525
Stein, Hall & Co., Inc. Hallmark
605 Third Avenue Jaguar
New York, NY 10016 Polyhall
Swift & Company Swift
Oakbrook, IL 60521
James Varley & Sons, Inc. Varco-Floc
1200 Switzen Avenue
St. Louis, MO 63147
W. E. Zimmie, Inc. Zimmite
810 Sharon Drive
Westlake, OH 44145
10.7.3 Other Chemicals
A large number of manufacturers generate sodium hydroxide; it is produced concurrently
with chlorine, and sometimes elemental sodium. The compound is produced at nearly
100 plant sites in this country, and bulk storage is available in most major cities. Designers
are referred to chemical suppliers in their geographical area for this commodity.
10-64
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Carbon dioxide is generated at some 75 locations around the country, and in addition
the liquefied gas is stored in bulk in nearly all major cities. Some twenty manufacturers
are involved in these operations, and designers are referred to local chemical suppliers
for further information on carbon dioxide.
About ten manufacturers produce soda ash in the United States. The 15 plant sites involved
are widely scattered geographically, as follows in Table 10-15:
TABLE 10-15
SODA ASH MANUFACTURERS
Ammonia-Soda Plants
Allied Chemical Corp., Syracuse, N.Y.; Detroit, Mich.; Baton Rouge,
La.
Diamond Alkali Co., Painesville, Ohio
Olin Mathieson Chemical Corp., Saltville, Va.; Lake Charles, La.
Pittsburgh Plate Glass Co., Barberton, Ohio; Corpus Christi, Texas
Wyandotte Chemicals, Corp., Wyandotte, Mich.
Natural Soda Plants
American Potash and Chemical Corp., Trona, Calif.
FMC Corp., Green River, Wyoming
Pittsburgh Plate Glass Co., Bartlett, Calif.
Stauffer Chemical Co., Westend, Calif.; Green River, Wyoming
Carbonation of Caustic Liquors
Dow Chemical Co., Freeport, Texas
10-65
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The preceding list is from a reference written to supply such information on these and
many other related compounds:
Faith, W. L.; Keyes, D. B.;
& Clark, R. L., Industrial
Chemicals, 3rd Ed., John
Wiley & Son, New York (1965).
10.8 Chemical Feeders
Chemical feed systems must be flexibly designed to provide for a high degree of reliability
in light of the many contingencies which may affect their operation. Thorough waste
characterization in terms of flow extremes and chemical requirements should precede the
design of the chemical feed system. The design of the chemical feed system must take
into account the form of each chemical desired for feeding, the particular physical and
chemical characteristics of the chemical, maximum waste flows and the reliability of the
feeding devices.
In suspended and colloidal solids removal from wastewaters the chemicals employed are
generally in liquid or solid form. Those in solid form are generally converted to solution
or slurry form prior to introduction to the wastewater stream; however, some chemicals
are fed in a dry form. In any case, some type of solids feeder is usually required. This
type of feeder has numerous different forms due to wide ranges in chemical characteristics,
feed rates and degree of accuracy required. Liquid feeding is somewhat more restrictive,
depending mainly on liquid volume and viscosity.
The capacity of a chemical feed system is an important consideration in both storage
and feeding. Storage capacity design must take into account the advantage of quantity
purchase versus the disadvantage of construction cost and chemical deterioration with time.
Potential delivery delays and chemical use rates are necessary factors in the total picture.
Storage tanks or bins for solid chemicals must be designed with proper consideration of
the angle of repose of the chemical and its necessary environmental requirements, such
as temperature and humidity. Size and slope of feeding lines are important along with
their materials of construction with respect to the corrosiveness of the chemicals.
Chemical feeders must accommodate the minimum and maximum feeding rates required.
Baker (12) indicates that manually controlled feeders have a common range of 20:1, but
this range can be increased to about 100:1 with dual control systems. Chemical feeder
control can be manual, automatically proportioned to flow, dependent on some form of
process feedback, or a combination of any two of these. More sophisticated control systems
are feasible if proper sensors are available. If manual control systems are specified with
the possibility of future automation, the feeders selected should be amenable to this
conversion with a minimum of expense. An example would be a feeder with an external
motor which could easily be replaced with a variable speed motor or drive when automation
10-66
-------
is installed (12). Standby or backup units should be included for each type of feeder
used. Reliability calculations will be necessary in larger plants with a greater multiplicity
of these units. Points of chemical addition and piping to them should be capable of
handling all possible changes in dosing patterns in order to have proper flexibility of
operation. Designed flexibility in hoppers, tanks, chemical feeders and solution lines is
the key to maximum benefits at least cost (12).
Liquid feeders are generally in the form of metering pumps or orifices. Usually these
metering pumps are of the positive-displacement variety, plunger or diaphragm type. The
choice of liquid feeder is highly dependent on the viscosity, corrosivity, solubility, suction
and discharge heads, and internal pressure-relief requirements (13). Some examples are
shown in Figure 10-13 and Figure 10-14. In some cases control valves and rotameters
may be all that is required. In other cases, such as lime slurry feeding, centrifugal pumps
with open impellers are used with appropriate controls (14). More complete descriptions
of liquid feeder requirements can be found in the literature and elsewhere (13).
Solids characteristics vary to a great degree and the choice of feeder must be considered
carefully, particularly in the smaller-sized facility where a single feeder may be used for
more than one chemical. Generally, provisions should be made to keep all chemicals
cool and dry. Dryness is very important, as hygroscopic (water absorbing) chemicals may
become lumpy, viscous or even rock hard; other chemicals with less affinity for water
may become sticky from moisture on the particulate surfaces, causing increased arching
in hoppers. In either case, moisture will affect the density of the chemical and may
result in under-feed. Dust removal equipment should be used at shoveling locations, bucket
elevators, hoppers and feeders for neatness, corrosion prevention and safety reasons.
Collected chemical dust may often be used.
The simplest method for feeding solid chemicals is by hand. Chemicals may be preweighed
or simply shoveled or poured by the bagful into a dissolving tank. This method is of
economic necessity limited to very small operations, or to chemicals used in very weak
solutions.
Because of the many factors, such as moisture content, different grades and compressibility,
which can affect chemical density (weight to volume ratio), volumetric feeding of solids
is normally restricted to smaller plants, specific types of chemicals which are reliably
constant in composition and low rates of feed. Within these restrictions, several volumetric
types are available. Accuracy of feed is usually limited to plus or minus 2% by weight
but may be as high as plus or minus 5%.
10-67
-------
FIGURE 10-13
PLUNGER TYPE METERING PUMP
(COURTESY OF WALLACE & TIERNAN)
/
DISCHARGE VALVE
PLUNGER
SUCTION VALVE
10-68
-------
FIGURE 10-14
DIAPHRAGM TYPE METERING PUMP
(COURTESY OF WALLACE & TIERNAN)
DISCHARGE VALVE
10-69
-------
One type of volumetric dry feeder uses a continuous belt of specific width moving from
under the hopper to the dissolving tank. A mechanical gate mechanism regulates the
depth of material on the belt, and the rate of feed is governed by the speed of the belt
and/or the height of the gate opening. The hopper normally is equipped with a vibratory
mechanism to reduce arching. This type of feeder is not suited for easily fluidized materials.
Another type employs a screw or helix from the bottom of the hopper through a tube
opening slightly larger than the diameter of the screw or helix. Rate of feed is governed
by the speed of screw or helix rotation. Some screw-type designs are self-cleaning, while
others are subject to clogging. Figure 10-15 shows a typical screw-feeder.
Most remaining types of volumetric feeders generally fall into the positive-displacement
category. All designs incorporate some form of moving cavity of a specific or variable
size. In operation, the chemical falls by gravity into the cavity and is more or less fully
enclosed and separated from the hopper's feed. The size of the cavity, and the rate
at which the cavity moves and is discharged, governs the amount of material fed. The
positive control of the chemical may place a low limit on rates of feed. One unique
design is the progressive cavity metering pump, a non-reciprocating type.
Positive-displacement feeders often utilize air injection to enhance flowability of the
material. Some examples of positive-displacement units are illustrated in Figure 10-16.
The basic drawback of volumetric feeder design, i.e., its inability to compensate for changes
in the density of materials, is overcome by modifying the volumetric design to include
a gravimetric or loss-in-weight controller. This modification allows for weighing of the
material as it is fed. The beam balance types measure the actual mass of material. This
is considerably more accurate, particularly over a long period of time, than the less common
spring-loaded gravimetric designs. Gravimetric feeders are used where feed accuracy of
about plus or minus 1% is required for economy, as in large-scale operations and for
materials which are used in small, precise quantities. It should be noted, however, that
even gravimetric feeders cannot compensate for weight added to the chemical by excess
moisture. Many volumetric feeders may be converted to loss-in-weight function by placing
the entire feeder on a platform scale which is tared to neutralize the weight of the feeder.
Good housekeeping and need for accurate feed rates dictate that the gravimetric feeder
be shut down and thoroughly cleaned on a regular basis. Although many of these feeders
have automatic or semi-automatic devices which compensate to some degree for
accumulated solids on the weighing mechanism, accuracy is affected, particularly on humid
days when hygroscopic materials are fed. In some cases, built-up chemicals can actually
jam the equipment.
10-70
-------
FIGURE 10-15
SCREW FEEDER
MOTOR AND
GEAR REDUCER
SOLUTION
CHAMBER
HOPPER
ROTATING &
RECIPROCATING
PEED SCREW
SOLUTION
LEVEL
JET MIXER
10-71
-------
FIGURE 10-16
POSITIVE DISPLACEMENT SOLID FEEDER-ROTARY
10-72
-------
No discussion of feeders is complete without at least passing reference to dissolvers, as
any metered material must be mixed with water to provide a chemical solution of desired
strength. Most feeders, regardless of type, discharge their material to a small dissolving
tank which is equipped with a nozzle system and/or mechanical agitator depending on
the solubility of the chemical being fed. Solid materials, such as polyelectrolytes, may
be carefully spread into a vortex spray or washdown jet of water immediately before
entering the dissolver. It is essential that the surface of each particle become thoroughly
wetted before entering the feed tank to ensure accurate dispersal and to avoid clumping,
settling or floating.
A dissolver for a dry chemical feeder is unlike a chemical feeding mechanism, which by
simple adjustment and change of speed can vary its output tenfold. The dissolver must
be designed for the job to be done. A dissolver suitable for a rate of 10 Ib/hr may
not be suitable for dissolving at a rate of 100 Ib/hr. As a general rule, dissolvers may
b oversized, but dissolvers for commercial ferric sulfate or lime slakers do not perform
well if greatly oversized.
It is essential that specifications for dry chemical feeders include specifications on dissolver
capacity. A number of factors need to be considered in designing dissolvers of proper
capacity. These include detention times and water requirements, as well as other factors
specific to individual chemicals.
The capacity of a dissolver is based on detention time, which is directly related to the
wettability of rate of solution of the chemical. Therefore, the dissolver must be large
enough to provide the necessary detention for both the chemical and the water at the
maximum rate of feed. At lower rates of feed, the strength of solution or suspension
leaving the dissolver will be less, but the detention time will be approximately the same
unless the water supply to the dissolver is reduced. When the water supply to any dissolver
is controlled for the purpose of forming a constant strength solution, mixing within the
dissolver must be accomplished by mechanical means, because sufficient power will not
be available from the mixing jets at low rates of flow. Hot water dissolvers are also
available in order to minimize the required tankage.
The foregoing descriptions give some indication of the wide variety of materials which
may be handled. Because of this variety, a modern facility may contain any number
of a variety of feeders with combined or multiple materials capability. Ancillary equipment
to the feeder also varied according to the material to be handled. Liquid feeders encompass
a limited number of design principles which account for density and viscosity ranges. Solids
feeders, relatively speaking, vary considerably due to the wide range of physical and
chemical characteristics, feed rates and the degree of precision and repeatability required.
10-73
-------
TABLE 10-16
TYPES OF CHEMICAL FEEDERS
Limitations
Type of feeder
Dry feeder:
Volumetric:
Oscillating plate
Oscillating throat (universal)
Rotating disc
Rotating cylinder (star)
Screw
Ribbon
Belt
Gravimetric:
Continuous-belt and scale
Loss in weight
Solution feeder:
Nonpositive displacement:
Decanter (lowering pipe)
Orifice
Rotameter (calibrated valve)
Loss in weight (tank with
control valve).
Positive displacement:
Rotating dipper
Proportioning pump:
Diaphragm
Piston
Gas feeders:
Solution feed
Direct feed
Use
Any material, granules or
powder.
Any material, any particle
size.
Most materials including
NaF, granules or powder.
Any material, granules or
powder.
Dry, free flowing material,
powder or granular.
Dry, free flowing material,
powder, granular, or lumps.
Dry, free flowing material up
to IVi-inch size, powder or
granular.
Dry, free flowing, granular
material, or floodable
material.
Most materials, powder
granular or lumps.
Most solutions or light
slurries.
Most solutions
Clear solutions
Most solutions
Most solutions or slurries
Most solutions. Special unit
for 5% slurries. 1
Most solutions, light
slurries.
Chlorine
Carbon dioxide
Chlorine
Capacity
(cubic feet per hr)
0.01 to 35
0.002 to 100
0.01 to 1.0
8 to 2,000
or
7.2 to 300
0.05 to 18
0.002 to 0.16
0.1 to 3,000
0.02 to 2
0.02 to 80
0.01 to 10
0.16 to 5
0.005 to 0.16
or
0.01 to 20
0.002 to 0.20
0.1 to 30
0.004 to 0.15
0.01 to 170
8,000 Ib/day max
6,000 Ib/day max
300 Ib/day max
Range
40 to 1
40 to 1
20tol
lOtol
or
100 to 1
20tol
lOtol
lOtol
to
100 to 1
100 to 1
100 to 1
100 to 1
10 to 1
lOtol
30tol
100 to 1
100 to 1
20tol
20tol
20tol
lOtol
Use special heads and valves for slurries.
10-74
-------
Table 10-16 describes several types of chemical feeders commonly used in wastewater
treatment.
10.9 Additional Reading
10.9.1 General
1. Water Quality and Treatment, Third Edition, American Water Works Association,
Inc., McGraw-Hill, NY (1971).
2. Water Treatment Plant Design, American Society of Civil Engineers, American
Water Works Association, and Conference of State Sanitary Engineers, NY
(1969).
3. Ockershausen, R. W., "Safe Handling of Water Works Chemicals," JAWWA, 63:6,
p. 336 (1971).
4. "Chemicals Used in Treatment of Water and Wastewater," BIF Reference No.
1.21-15, BIF Industries, Providence, RI (1970).
10.9.2 Aluminum Compounds
5. Truitt, V., "We Converted to Liquid Alum Treatment," Public Works, 99:11,
p. 88 (1968).
6. "Aluminum Sulfate," Allied Chemical Corporation, Morristown, NJ (1968).
7. "Alum," American Cyanamid Company, Wayne, NJ (1964).
8. "Aluminum Sulfate," Stauffer Chemical Company, NY (1967).
10.9.3 Iron Compounds
9. "Handling Ferric Chloride," Dow Chemical Company, Midland, MI (1965).
10. "Pennsalt Ferric Chloride," Pennsalt Chemicals Corporation, Philadelphia, PA
(1965).
10-75
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10.9.4 Lime
11. "Chemical Lime Facts," National Lime Association, Washington, D.C. (1964).
12. Carr, R. L., "Limestone to Lime to Slaked Lime," Water and Sewage Works,
Reference No. 113:R.N., p. R-61 (1966).
13. "Chemicals Used in Water and Wastewater Treatment - Calcium Hydroxide
(Hydrated Lime)," Water and Wastes Engineering, 7:1, p. 53 (1970).
14. "Chemicals Used in Water and Wastewater Treatment - Quicklime," Water and
Wastes Engineering, 7:5, p. 52 (1970).
15. "Lime Handling, Application & Storage in Treatment Processes," National Lime
Association, Washington, D. C. (1971).
16. "Chemistry and Technology of Lime and Limestone," Boynton, R. S., John
Wiley & Sons, New York, NY (1966).
10.9.5 Soda Ash
17. "Soda Ash," Allied Chemical Corporation, NY (1966).
18. "What's Different About FMC Soda Ash?", FMC Corporation, NY (1964).
19. "Stauffer 'Natural' Soda Ash," Stauffer Chemical Company, NY (1964).
20. "Wyandotte Soda Ash," Wyandotte Chemicals Corporation, Wyandotte, Michigan
(1955).
10.9.6 Caustic Soda
21. "Caustic Soda," PPG Industries, Inc., Chemical Division, Pittsburgh, PA (1969).
22. "Unloading Liquid Caustic from Tank Cars," Manufacturing Chemists
Association, Washington, D.C. (1968).
23. "Properties and Essential Information for Safe Handling and Use of Caustic
Soda," Manufacturing Chemists Association, Washington, D.C. (1968).
10-76
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24. "Caustic Soda," Allied Chemical Corporation, NY (1963).
25. "Caustic Soda Handbook," Diamond Alkali Company, Cleveland, OH (1967).
26. "Dow Caustic Soda Handbook," Dow Chemical Company, Midland, MI
(Undated).
27. "Hooker Caustic Soda," Hooker Chemical Corporation, Niagara Falls, NY (1966).
28. "Olin Caustic Soda," Olin Mathieson Chemical Corporation, NY (1961).
29. "Pennsalt Caustic Soda," Pennsalt Chemicals Corporation, Philadelphia, PA
(1964).
30. "Caustic Soda," Stauffer Chemical Company, NY (1965).
31. "Caustic Soda by Wyandotte," Wyandotte Chemicals Corporation, Wyandotte,
MI (1961).
10.10 References
1. "Chemicals Used in Water and Wastewater Treatment-Aluminum Sulfate," Water
and Wastes Engineering, Dec. 6:12, p. 46 (1969).
2. "Chemicals Used in Water and Wastewater Treatment-Ferric Chloride," Water
and Wastes Engineering, 7:3, p. 65 (1970).
3. "Ferri-Floc for Water and Wastewater Treatment," Cities Service Co., Atlanta,
GA (1972).
4. Schworm, W. B., "Iron Salts for Water and Waste Treatment," Public Works,
94:10, p. 118 (1963).
5. "Lime for Water and Wastewater Treatment," BIF Reference No. 1.21-24, BIF
Industries, Providence, RI (June, 1969).
6. National Lime Association, direct communication (May, 1971).
7. "Caustic Soda," PPG Industries, Inc., Chemical Division, Pittsburgh, PA (1969).
10-77
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8. Haney, P. D. and Hamann, C. L., "Recarbonatkm and Liquid Carbon Dioxide,"
JAWWA, 61:10, p. 512 (1969).
9. Gulp, R. L. and Gulp, G. L., Advanced Wastewater Treatment, Van
Nostrand-Reinhold Company, NY (1971).
10. Carr, R. L., "Polyelectrolyte Coagulant Aids - Dry and Liquid Handling and
Application," Water and Sewage Works, Reference No. 114:R.N., p. 4-64 (1967).
11. Russo, F. and Carr, R. L., "Polyelectrolyte Coagulant Aids and
Flocculants: Dry and Liquid, Handling and Application," Water and Sewage
Works, Reference No. 117:R.N., p. R-72 (1970).
12. Baker, R. J., "Chemical Feed Systems Determine Plant Efficiency and
Reliability," Water and Sewage Works, Reference No. 116:R-21 (November,
1969).
13. Russo, F., and Carr, R. L., Jr., "Polyelectrolyte Coagulant Aids and
Flocculants: Dry and Liquid, Handling and Application," Water and Sewage
Works, Reference No. 117:R-72 (November, 1970).
14. Gulp, R. L., and Gulp, G. L., Advanced Wastewatewater Treatment, Van Nostrand
Reinhold Company, NY (1971).
15. R. P. Lowe, "Chemical Feed Systems," 10th Annual Water Conf. of Eng. Soc.
of W. PA (October 17-19, 1949).
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CHAPTER 11
SLUDGE HANDLING AND DISPOSAL
11.1 Introduction
A high proportion of wastewater treatment capital and operating cost is due to sludge
handling and disposal. Consequently, it is important that design engineers have reliable
information on quantities and dewatering properties of sludges so that solids handling
facilities can be properly designed.
Information on sludges produced in conventional wastewater treatment is widely scattered,
and results from different locations are not related well to each other. There is even
less information available on sludges produced by the newer processes for phosphorus
removal.
There has recently been a concerted effort to provide this necessary information. One
group involved is the staff of the Ultimate Disposal Research Program operating as part
of the Advanced Waste Treatment Research Laboratory of the National Environmental
Research Center in Cincinnati.
Material in this chapter comes from information compiled and organized by that group.
Their basic data were obtained in several ways:
a. Grants and contracts aimed at obtaining comparative information
on phosphorus-removal sludges have been funded.
b. Reports have been prepared which summarize information
obtained on sludges in EPA Research and Development Grants
where useful sludge data have been collected.
c. Field crews have visited certain plants where promising processes
are being implemented and have obtained comparative data on
sludges involved.
d. Visits have been made to municipalities and sludge data have been
obtained from plant staffs.
e. Published literature has been surveyed and information of value
on sludges has been extracted.
f. Memoranda have been prepared to organize the information
available and to respond to specific needs.
11-1
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g. Experimental investigations of sludge dewatering are being
conducted at several locations by Advanced Waste Treatment
Laboratory staff.
All of these routes are still being actively pursued, although the collection of information
is far from complete. However, there was such urgent need for this information
(e.g., inclusion in this manual) that an interim report was prepared, and this chapter was
developed from it.
Progress on sludge handling and disposal technology will be dealt with in detail in a
forthcoming design manual on the subject (1).
11.2 Generalized Assessment of Filtration Rates
At least two groups (2) (3) have developed cost information computer programs which
can be used to predict the cost of alternative processing schemes for wastewater treatment
and for phosphorus removal processes. These programs have included in their inherent
logic (not as input data) typical values for dewatering rates, normal sludge mass, and
additional sludge produced by phosphorus removal processes.
Aside from the output from these cost estimated computer programs, reliable estimates
of the dewaterability of various sludges can be of substantial value to those concerned
with process design. That type of information is presented in this chapter.
When attempting to decide on proper values for sludge dewatering properties, there is
a statistical problem of within-plant variations and between-plant variations. The
between-plant variations are sometimes extreme. This is illustrated in the following
hypothetical compilation:
TABLE 11-1
HYPOTHETICAL COMPILATION OF DEWATERING RATES
Vacuum Filter Yield (Ib/hr-sq ft)
Avg. (40 Plants) Std. Deviation
Digested Mixed Sludge 4.0 2.0
Raw Mixed Sludge 5.0 3.0
11-2
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The designer wishing to use data of this type to predict within-plant differences between
raw and digested sludges will find tabulated values are of little value because standard
deviations are so great. Within-plant variations between raw and digested sludge are less
than indicated in the table. Unfortunately, within-plant comparisons are virtually
unavailable because usually only one type of sludge at a given plant is dewatered.
Information relating various within-plant sludges to each other should eventually be
collected. In this chapter there has been an attempt to estimate such relationships based
on experience and information from the sources previously cited. Relative values are
suggested here. These generalized figures should be used only when there is no other
information available.
The information in this section relates only to vacuum filtration. The important area
of centrifugation has been introduced in other chapters, when such information was
available. Section 11.4 also describes some case studies in which centrifugation was used
as a dewatering means, but data are insufficient for development of generalized design
parameters.
There has been no consideration given to lime sludges in this introductory section. Sludges
produced by lime addition, even when added to the primary clarifier, contain large
quantities of calcium carbonate and filter better than conventional or alum and iron sludges.
In addition, a different approach is needed than for alum, iron, and conventional sludges.
Section 11.4 describes ease studies where lime sludge is dewatered by various means, and
Section 11.3.2 presents fundamentals involved.
11.2.1 Conventional Sludges
The digested total sludge from a conventional primary plus activated sludge plant is
probably the commonest sludge encountered in wastewater treatment. Consequently, it
has been selected as the standard against which other sludges are compared.
Table 11-2 shows estimated relative filtration yields and conditioning costs of various
wastewater sludges compared to a standard digested primary plus waste activated sludge.
These values are for belt-type vacuum filters operated at optimum conditions, preceded
by a reasonable processing sequence including thickening if clarifier sludge is dilute. Filter
cloth selection was appropriate for the situation to get satisfactory solids capture.
11-3
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TABLE 11-2
DESIGN FACTORS FOR FILTRATION OF CONVENTIONAL SLUDGES
Relative* Relative Cost Cake Solids
Sludge Type Yield of Conditioning** Content
Raw Primary
Plus WAS 1.4 0.7 28
Digested Primary
Plus WAS 1.0 1.0 26
Raw Primary
Plus TF 1.8 0.6 32
Digested Primary
Plus TF 1.3 0.9 30
Digested Primary 2.3 0.7 33
Raw Primary 3.2 0.5 35
*Avg. yield for digested primary plus WAS is estimated at 4.0 Ib/hr-sq ft
**Average cost of conditioning is $10/dry ton, 1974 costs
An average value of 4.0 Ib/hr-sq ft is suggested for the yield when filtering digested primary
plus waste activated sludge. If a good estimate of the yield for this or any other sludge
is available, it should be utilized and estimates of the yields for the other sludges obtained
by means of the factors in Table 11-2.
Mass of sludge produced in wastewater treatment can be estimated by means of
well-established procedures discussed in Section 11.3. Reduction in mass by digestion
can be estimated by volatile solids reduction. Since digestion reduces mass approximately
50 percent, required filter area will be less for digested sludge than for raw sludge despite
slightly higher yields for raw sludges.
Some shortcomings of the tabulated numbers are evident. No influence of solids content
is indicated, although it is well known that, up to about 8% solids, filter yield is
proportional to solids content. It must be presumed that appropriately designed thickeners
are used. It becomes clear that when a filter yield is chosen, there are implicit assumptions
about the operations upstream.
Effect of elutriation is not included in Table 11-2. Elutriation would reduce lime demand
substantially. Polymer required for conditioning would be reduced as well; elutriation
11-4
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is virtually a necessity for some digested sludges (4).
No indication of range in solids content of filtered sludge cake is given. This is of
importance when the sludge is to be incinerated. Some data (5) indicate that digested
primary and waste activated sludge would have a solids content of about 28%. For raw
primary and waste activated sludge, it would be about 30%. Slightly lower values are
suggested in Table 11-2.
11.2.2
Iron and Alum Sludges
Sludges produced when adding ferric or aluminum salts have a substantially different
character than conventional sludges. Quantities are greater and handling and dewatering
properties are different. Changing the addition point in the processing scheme also produces
changes in the character of the sludge.
Table 11-3 compares estimated vacuum filtration yields and conditioning costs of the
various metal-effected sludges relative to conventional digested primary plus waste activated
sludge.
TABLE 11-3
DESIGN FACTORS FOR FILTRATION
OF CONVENTIONAL PLUS METAL SLUDGES
Sludge Type
Aluminum to Primary
Raw Primary Plus WAS
Digested Primary Plus WAS
Raw Primary Plus TF
Digested Primary Plus TF
Digested Primary
Raw Primary
Aluminum to Aerator
Relative
Yield*
1.2
0.85
1.35
1.0
1.05
1.5
Relative Cost
of Conditioning*
1.3
1.4
1.2
1.3
1.0
0.9
Raw Primary Plus WAS
Digested Primary Plus WAS
1.2
0.85
1.3
1.4
11-5
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Iron to Primary
Raw Primary Plus WAS 1.3 1.1
Digested Primary Plus WAS 0.95 1.2
Raw Primary Plus TF 1.5 1.0
Digested Primary Plus TF 1.1 1.1
Digested Primary 1.2 0.8
Raw Primary 1.6 0.7
Iron to Aerator
Raw Primary Plus WAS 1.3 1.1
Digested Primary Plus WAS 0.95 1.2
*Yields and costs are related to yield and cost obtained with digested primary plus WAS
when no chemicals are added to the wastewater. See Table 11-2.
Estimates will be more accurate if absolute yields for one or more sludges can be determined
for a particular situation. The factors in the table can then be used to predict other
yields. If no information is available, filter yield for digested conventional primary plus
waste activated sludge may be estimated to be 4.0 Ib/hr-sq ft.
11.3 Quantity of Phosphorus Removal Sludges
11.3.1 Iron and Alum Sludges
When ferric iron or trivalent aluminum are added to wastewater containing phosphorus,
a metal phosphate precipitates and excess metal forms an insoluble hydroxide. The
hydroxide acts as a flocculant precipitate which carries with it other suspended solids
when it settles. In addition, soluble organic material apparently is adsorbed on the
precipitate and removed, as discussed later.
The amount and nature of additional sludge formed is a function of the process addition
point. For example, addition to the primary clarifier causes removal of much more
suspended solids than occurs in conventional sedimentation. Also, soluble organic material
is removed there. However, the biological stage which follows receives a lower BOD
wastewater; consequently, the amount of biological sludge is reduced correspondingly.
There are available tabulations which present average observed values of the amount of
sludge typically produced when iron and aluminum are added to wastewater (6). However,
there is still need for a more precise method of predicting sludge production so designers
can more accurately size sludge processing equipment.
11-6
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A detailed procedure is presented here for the case where these metals are added to the
conventional processing sequence of primary treatment plus either activated sludge or
trickling filter. In this procedure the chemicals may be added in primary or secondary
treatment, or as a tertiary process. The method requires that the following information
be available: complete processing sequence, wastewater characteristics (e.g., SS, BOD,
soluble TOC), and efficiency of solids removal in clarifiers. Methods are outlined for
predicting the following: chemical sludge produced, extra sludge from improved clarifier
efficiency, extra sludge from removal of soluble materials by the chemicals, and reduced
amount of biological sludge when biological plant load is reduced. In this manner, total
sludge production, related to processing conditions and wastewater quality, can be
determined. The method is given in Sections 11.3.1.1 through 11.3.1.5.
11.3.1.1
Theoretical Production of Chemical Sludge
Chemical sludge production may be calculated by assuming the metal added reacts first
with phosphorus compounds which are thereby removed, and that any excess metal forms
hydroxide. The calculation is illustrated in Table 11-4.
TABLE 11-4
CALCULATION OF CHEMICAL SLUDGE PRODUCTION
Given: Pin =10 mg/1
•out
= 1 mg/1
Al Case:
Fe Case:
dose of 19 mg/1 Al
(Al/P atomic ratio = 2.18)
dose of 40 mg/1 Fe
(Fe/P atomic ratio = 2.22)
Aluminum
Stoichiometry:
Al + PO4 = Al PO4
Al + 3 OH = Al (OH)3
11-7
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(9 mg/1 P)/(31) = 0.29 millimoles/1 Al PO4
(M.W. = 122)
(19 mg/1 Al)/27 = 0.70 mmol/1 Al
0.70 - 0.29 = 0.41 mmol/1 Al excess, goes to Al (OH)3
(MW = 78)
(0.29 mmol/1 Al PO4)(122) = 35.4 mg/1 Al PO4
(0.41 mmol/1 Al (OH)3)(78) = 32.0 mg/1 Al (OH)3
Total = 35.4 + 32.0 = 67.4 mg/1
Sludge Produced:
67.4 mg chemical sludge per liter of wastewater treated
or 67.4/19 = 3.55 mg chemical sludge per mg Al
or 67.4/9 = 7.48 mg chemical sludge per mg P removed
Iron
Stoichiometry:
Fe + PO4 = Fe PO4
Fe + 3 OH = Fe (OH)3
(9 mg/1 P)/(31) = 0.29 mmol/1 Fe PO4
(M.W. = 150.8)
(40 mg/1 Fe)/(55.84) = 0.72 mmol/1 Fe
0.72 - 0.29 = 0.43 mmol/1 Fe excess, goes to Fe(OH)3
(M.W. = 106.8)
(0.29 mmol/1 Fe PO4)( 150.8) = 43.7 mg/1 Fe PO4
(0.43 mmol/1 Fe (OH)3)( 106.8) = 45.9 mg/1 Fe (OH)2
Total = 43.7 + 45.9 = 89.6 mg/1
11-8
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Sludge Produced:
89.6 mg of chemical sludge per liter of wastewater treated
or 89.6/40 = 2.13 mg chemical sludge per mg Fe
or 89.6/9 = 9.96 mg chemical sludge per mg P removed
The illustration in Table 11-4 shows that at approximately the same atomic ratio of metal
to phosphorus, 48 percent more chemical sludge is predicted for iron addition than for
aluminum addition.
Sludge production is related in Table 11-4 to mass of phosphorus removed per liter of
sewage treated. This is the most rational procedure when sludge mass is calculated.
However, it is frequently convenient to relate sludge mass to the mass of metal added.
The following are estimated values for the ratio of sludge mass to metal dose as the metal/P
atomic ratio increases from zero to infinity:
Al: 4.52 to 2.89 mg sludge/mg Al
Fe: 2.70 to 1.92 mg sludge/mg Fe
Information to verify these calculated values is scarce. Swedish experience (7) indicates
the additional weight of sludge may be calculated from the following relationships:
sludge (mg/1 of wastewater) = (4) (Al dose, mg/1)
sludge (mg/1 of wastewater) = (2.5) (Fe dose, mg/1)
That reference does not elaborate on the source of information although it was drawn
from Swedish practice at the time. It is not stated whether additional mass from improved
suspended and dissolved solids removal is included. The great majority of plants in Sweden
were adding chemicals as a post-treatment after biological treatment. Consequently, it
may be presumed that these figures indicate total increase in sludge mass from Al or
Fe added to the final effluent from a biological treatment plant.
Other studies (8) estimated additional sludge produced by alum addition to the activated
sludge process. They report 4 to 3 mg addition solids per mg of Al as dose level increased
from 10 to 20 mg Al/1.
Another report (9) has estimated chemical sludge production when alum was added to
secondary effluents. The following relationship was observed:
chemical sludge (mg/1 of wastewater) = (5.2)(A1 dose, mg/1)
11-9
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Al/P atomic ratio range: 2.0 to 2.8
This is about 20 percent higher than the Swedish figure and about 45 percent higher
than indicated in Table 11-4 for the stoichiometric relationship.
Data available, then, indicate chemical sludge mass higher than the values represented by
the simple stoichiometry indicated in Table 11-4. There is reason to conclude that
chemical sludge production could be higher than stoichiometric values because other cations
(such as heavy metals) and anions can be incorporated into the complex ligand structures
of aluminum and ferric hydroxides. Consequently, it is recommended that with either
Al or Fe, sludge mass be estimated to be 35 percent higher than the simple stoichiometric
values.
11.3.1.2 Theoretical Production of Biological Sludge
Biological sludge is produced and wasted in both the activated sludge process and the
trickling filter process.
One source (10) reports the following equation for excess sludge production in the
conventional activated sludge process:
Wy/M = (0.79 F/M) - (0.0714)
where Wy = Ib volatile suspended solids produced per day
M = Ib mixed liquor volatile suspended solids in the aerator
F = Ib 5-day BOD removed per day
Rearranging this equation,
Wy/F = (0.79) - (0.0714/F/M)
The equation now gives the mass of volatile solids produced per unit mass of BOD removed
(WV/F) as a function of the food to mass ratio.
A recent report (11) states that 50 percent of the sum of the BOD and suspended solids
fed to a trickling filter is converted to waste sludge.
If suspended solids and BOD concentrations in the wastewater as it enters the biological
treatment step are known, sludge production can be estimated by means of the above
relationships.
11-10
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11.3.1.3 Increase in Production of Sludge From Sewage Solids
Suspended solids and dissolved solids in wastewater are removed by chemical treatment.
Sludge mass is increased by the direct contribution of the mass of these removed solids.
One report (12) has summarized data on suspended solids removal from well designed
physical-chemical pilot plants. For an average sewage and effective clarifiers, suspended
solids removal of 85 percent can be achieved. If the sewage is high in suspended solids,
removals will be better. This figure is substantially higher than the 50-75% removal range
suggested previously (13). However, those earlier figures included plants where chemicals
were added to an existing facility which might have been overloaded or were functioning
poorly for other reasons.
The mass of sludge produced from suspended solids removed is calculated directly by
multiplying the change in suspended solids concentration by the volume of sewage.
There is also evidence (12) which indicates an appreciable mass of dissolved materials
(i.e., filtrable through a 0.45 micron membrane filter) is removed by chemical treatment.
The following table presents this information along with data collected from other sources.
TABLE 11-5
REMOVAL OF DISSOLVED MATERIALS DURING CHEMICAL ADDITION
Source Percent Removals
_Fe_ Al_
SCOD STOC SCOD STOC
Eimco
pilot plant (14) 40 30
laboratory (14) ca. 30
Friedman
et al (15) 30
Farrell (16) 27
Westrick and
Cohen (17) 45 42
11-11
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Note: SCOD is soluble COD
STOC is soluble TOC
Some of that date (16) represent the average removal obtained in seven batch experiments
in which raw sewage was treated with alum in a small clarifier (150 liter batches). Al/P
atomic* ratio ranged from 1.6 to 2.8 and removals from 20 to 32 percent.
Based on the above data, it is concluded that soluble TOC removals will average 30 percent.
Data are scant for COD and none appear available for BOD. It is assumed, however,
that removals will be the same for STOC.
The mass of sludge associated with the soluble TOC removed is more than the carbon
content involved. It is estimated that, on the average, the soluble compounds removed
by Al and Fe treatment will average 40 percent carbon on a volatile solids basis. This
figure is in accord with the carbon content of compounds tested in developing the TOC
test for wastewaters (18). Thus, to obtain an estimate of the mass of sludge (volatile
basis) produced, TOC removed should be multiplied by 1/0.40 or 2.5.
If only soluble COD removed or soluble BOD removed is known, the problem is more
difficult, primarily because the relationship between mass and COD or BOD changes as
the degree of oxidation of the sewage changes. For a secondary effluent, volatile mass/COD
is generally considered to be about 1.5 (19). Assuming the BOD is about 70 percent
of the COD, volatile mass/BOD is 2.15. For raw sewage, which has a low degree of
oxidation, the COD is expected to be higher for a given mass. Volatile mass/COD for
a raw sewage is estimated to be 1.1, approximately 10 percent lower than lowest values
previously found (19) for secondary effluent. Volatile mass/BOD would be about 1.6.
In the above discussion, mass relates to mass of volatile solids. To obtain total mass,
non-volatile solids must be included. It is estimated that the ash content of these soluble
solids is about 15 percent. Mass of volatile solids should be multiplied by 100/(100-15)
or 1.18 to get total solids.
Calculation of sludge mass resulting from removal of soluble material will require
information on the absolute value of soluble TOC, COD or BOD. These values are rarely
measured. An examination of previous data (14), where both SCOD and STOC were
measured, showed that ratios of soluble to total COD and TOC in the influent sewage
were 0.30 and 0.38 respectively. Other data (16) showed an STOC/TOC ratio of 0.64.
If data on soluble values of TOC or COD are not available, soluble TOC or COD may
be estimated to be 50 percent of the total TOC or COD.
The correlations for predicting sludge production require information on total BOD
removal. Some data (12) indicate total BOD, COD and TOC reductions are about the
same and range from 50 to 85 percent depending on clarifier efficiency. Removal of
11-12
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65 percent is thereby chosen as a representative figure and will be used in example
calculations.
11.3.1.4 Summary of Calculation Procedures
The calculation procedure is described for various chemical addition points and is illustrated
with a specific example in Section 11.3.1.5. Factors involved in that example are
summarized here.
Chemicals Added to Primary
Chemical sludge is calculated on the basis of the phosphorus level and the chemical dose
in the manner illustrated in Table 4. Estimated mass should be increased by about
35 percent (see Section 11.3.1.1).
Additional suspended solids removals in primary treatment are estimated using quantities
derived per Section 11.3.1.3. The additional suspended solids mass removed should be
added directly to total sludge mass.
Mass resulting from dissolved solids removal has to be estimated indirectly from soluble
TOC, COD or BOD removal. Factors to be used in predicting mass have been discussed
in Section 11.3.1.3. Equations relating soluble TOC, COD and BOD input levels to mass
of soluble solids removed are summarized below:
(soluble solids mg/1) = STOCm(mg/l)x0.30x2.5xl.l8
= SCODin(mg/l)x0.30x l.lxl. 18
The last two equations apply only to influents before oxidative processes such as the
activated sludge process.
Estimates of total TOC, COD or BOD removals are needed to estimate the loading on
the biological processes which follow chemical treatment. However, this information is
not needed for estimating sludge formation in the chemical treatment step.
Chemicals Added to Activated Sludge
The chemical sludge is calculated in the manner described in Section 11.3.1.1.
11-13
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Addition of chemicals to the activated sludge process may produce a significant increase
in BOD and suspended solids removals. This is particularly true of a high rate activated
sludge plant or an overloaded plant. However, for a well-designed standard rate plant
addition of chemicals will produce little additional sludge other than the chemical sludge
itself. It is recommended that for a standard rate plant no increase in mass be included
for additional suspended solids. For special cases such as an existing high rate plant,
judgment will have to be made based on available information.
Chemicals Added After Biological Treatment
The chemical sludge is calculated as in Section 11.3.1.1. If the tertiary treatment follows
a standard rate activated sludge plant, additional sludge from improved suspended solids
and dissolved BOD removal can be neglected. For a trickling filter plant, however,
additional mass removed could be important. For one plant (20) a substantial
improvement in BOD and solids removal in a high rate trickling filter plant resulted from
alum addition before the final clarifier. In one series of runs, overall BOD removals
increased from 73% to 91% and suspended solids removal increased from 70% to 84%.
At another plant (21), under conditions where the final clarifier was not hydraulically
overloaded, alum added after a standard rate trickling filter brought BOD and suspended
solids from the 15-20 mg/1 levels to below 10 mg/1.
These results can be used to calculate additional sludge brought down with the chemical
sludge.
11.3.1.5 Illustrative Calculation
Sludge mass is calculated here for a plant utilizing primary treatment with aluminum
introduced before clarification, followed by standard rate activated sludge. Details of
the calculation are presented in Table 11-6. The existing plant is a properly designed
plant operating within its design limits, which is changed only by introduction of a means
to inject the chemical and flocculate it before primary clarification.
TABLE 11-6
CALCULATION OF TOTAL SLUDGE MASS WITH METAL ADDITION
Given: Plant utilizes primary clarification followed by standard rate activated sludge.
Alum is to be added in a flocculation tank prior to the primary clarifier. Estimate
the additional sludge produced by alum addition.
11-14
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Primary clarifier of the existing plant is designed for 50% suspended solids
removal (without chemical). Activated sludge plant is designed for an F/MLVSS
ratio of 0.4.
Overall phosphorus removal is to be 90%. Estimate that this can be achieved
at an Al/P atomic ratio of 2.0.
Influent to Primary: P 10 rng/1
Total BOD = 172 mg/1
Suspended Solids = 240 mg/1
Soluble TOC = 40 mg/1
Sludge from Primary Treatment - No Chemicals
Estimate 50% SS removal, 30% total BOD removal and 0% STOC removal.
240 x 0.50 - 120 mg/1
Sludge from Primary Treatment - Alum Added
Estimate 90% P removal, 85% SS removal, 65% total BOD removal and 30%
STOC removal.
Al Dose: (10/3I)(2)(27) = 17.4 mg/1
10 x 0.90 = 9.0 mg/1 P removal
9/31 = 0.29 meg (milliequivalent)/1 Al PO4
17.4/27 = 0.64 meg/1 Al
0.64 - 0.29 = 0.35 meg/1 Al excess
0.29 x 122 = 35.4 mg/1 Al PO4
0.35 x 78 = 27.3 mg/1 Al (OH)3
35.4 + 27.3 = 62.7 mg/1 sludge produced
Chemical Solids: Increase by 35%, per Section 11.3.1.1
62.7 x 1.35 = 84.5 mg/1
11-15
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Suspended Solids: 240 x 0.85 = 204 mg/1
Dissolved Solids: 40xO.30x2.5xl.18 = 35.4 mg/1
Total Sludge: 323.9 mg/1
Sludge from Activated Sludge - No Chemicals
Equation from Section 11.3.1.2
Wy/F = (0.79) - (0.0714/F/M)
F/M is given as 0.4
Wy/F = 0.79 - 0.18 = 0.61 Ib VSS/day/lb BOD removed per day
or 0.61 mg VSS/mg BOD removed
BOD in primary effluent is:
172 x 0.70 = 120 mg/1
BOD in final effluent is estimated to be almost 12 mg/1. See reference (6).
BOD removed = 120 - 12 = 108 mg/1
Therefore, 108 x 0.61 = 66 mg/1 VSS produced
Assuming the sludge is 70% volatile,
66/0.70 = 94 mg/1 waste sludge
Sludge from Activated Sludge - Al to Primary
BOD in primary effluent is:
172 x 0.35 = 60.2 mg/1
This is a very dilute waste. The existing activated sludge plant will be greatly
oversized. Air requirements and the contact time may be reduced.
Keeping F/M the same (0.4),
WV/F = 0.61 (see above)
11-16
-------
Estimate BOD in final effluent to be 12 mg/1 (see above).
BOD removed = 60.2 - 12 = 38.2 mg/1
Therefore, 38.2 x 0.61 = 23.3 mg/1 VSS produced
Assuming sludge is 70% volatile,
23.3/0.70 = 33.3 mg/1 waste sludge
Summary of Sludge Quantities Calculated
Sludge Quantities (mg/1)
No Chemical Al Added
Primary Clarifier
Sludge from SS 120 204
Sludge from Dissolved Solids 0 35.4
Chemical Sludge 0 84.5
Total Clarifier Sludge 120 323.9
Secondary Clarifier
Waste Activated Sludge 94 33.3
Total Sludge 214 357.2
Total Sludge Excepting
Chemical Sludge 214 272.7
It is clear that chemical addition produces a substantial increase in the total sludge mass.
The increase is greater than just the amount of chemical sludge because, in the no-chemical
case, there is substantially more conversion of organic substances to carbon dioxide and
water.
For a plant in which the chemical is added to the activated sludge process, the increase
in mass will be virtually all due to the chemical sludge, since essentially the same amount
of biological degradation is occurring as in the no-chemical case.
11-17
-------
In the illustration, total sludge production (mass) was increased 67 percent when adding
aluminum. The chemical sludge increased sludge mass 40 percent. The additional
27 percent sludge came from organic material which would have been oxidized had it
reached the activated sludge aerator. If the chemical had been added to the aerator,
sludge mass would have only increased 40 percent. The greater sludge mass which results
when the chemical is added to the primary must be balanced against the reduced dimensions
of the activated sludge aerator, the reduction in quantity of air needed, any differences
in the dewatering properties of the different sludges, and differences in chemical demand
which may be dependent on point of addition.
11.3.2 Lime Sludges
Lime is utilized for phosphorus removal in wastewater treatment by adding it just prior
to primary clarification or after secondary clarification as a tertiary treatment process.
As in the case of metal addition to wastewater, average values for the amount of sludge
produced have been tabulated (6). However, there is need for more precise information
applied to specific situations.
In response to this need, a procedure has been developed for predicting the mass of sludge
produced when wastewater is treated with lime in a tertiary process. Details of this
procedure are presented in this section. The method may be applied with minor
modification when lime is added just prior to the primary clarifier.
Unlike metal addition, the amount of sludge produced by lime addition depends not only
on the degree of phosphate removal desired, but also on alkalinity and other characteristics
of the wastewater. In fact, both the amount of sludge produced and the degree of
phosphorus removal achieved by raising the pH to a given level both depend on the nature
of the water.
The procedure in predicting sludge quantity requires first that pH level be estimated for
the desired phosphorus removal, then lime dose is estimated, and finally sludge mass can
be calculated. The individual steps are listed below:
a. From desired P-residual and knowledge of the water type, estimate
pH level.
b. From pH level and alkalinity, estimate lime dose.
c. From amount of P removed, calculate the hydroxyapatite sludge
produced.
d. From magnesium removal, calculate magnesium hydroxide sludge
produced.
11-18
-------
e. From calcium balance, calculate calcium carbonate sludge
produced.
f. From pH and alkalinity, verify that sufficient carbonate is present
to supply carbonate in sludge. If not, correct outlet calcium
estimate.
g. Include in total sludge production the inerts introduced with the
lime.
If lime is added prior to the primary settler, the procedure for estimating sludge for tertiary
lime addition may be used with some modification. However, the correlation between
pH and the ratio of lime dose to alkalinity was developed from data obtained by adding
lime to secondary effluents. There is no assurance it would fully apply to raw wastewater.
It would thus be desirable to establish lime dose needed by actual tests.
When lime is added prior to the primary settler, efficiency of the clarifier is increased
substantially. From a previous review (12), it is estimated that for well designed plants,
suspended solids removal efficiencies are in the order of 85 percent. This figure may
be used to estimate additional sludge solids removed in the primary clarifier by lime
addition.
The removal of dissolved organic solids by addition of lime prior to the primary appears
to be less than obtained with metals. One study (14) reported very small or negative
removals. Others (17) reported substantially lower removals with lime than with metal.
Considering the relatively large mass of sludge from lime treatment, the small amount
of dissolved organic solids brought down in primary treatment can be neglected.
Some prior data (14) have been utilized to estimate whether material balance calculations
similar to those described above correlate with measured sludge production. This case
is discussed in Section 11.4.2.2. Measured values were about 20 percent higher than the
material balance figures. It is probably advisable to utilize a factor of safety of this
magnitude in estimating chemical sludge production.
11.3.2.1 Estimation of Lime Dose and Sludge Production in Tertiary
Treatment
A quantitative calculation procedure for estimating the required lime dose and the quantity
of sludge produced in tertiary treatment of wastewater with lime is presented. On-site
testing for obtaining this information is recommended in specific plant designs.
11-19
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11.3.2.2 Lime Dose
The extent of phosphate removal is a function of a number of variables including chemical
composition of the waste water and the nature of the phosphate compounds contained.
The most important variable controlling phosphate removal and quality of effluent is pH
rather than quantity of lime added. Figure 11-1, reference (22), shows a correlation of
P-removal as a function of pH, utilizing data from several sources.
Another factor of importance is the hardness level of the wastewater. In some cases (23)
it was necessary to go to pH in excess of 11 to get good clarification with the low hardness
water. Others (24) obtained satisfactory clarification at pH of about 10 in the hard water.
The chemistry of lime-softening of natural waters is sufficiently well developed that the
charge of lime required to achieve a given removal of calcium and magnesium can be
predicted and the final pH can be closely estimated. With wastewaters, suspended and
dissolved organic materials as well as ortho and polyphosphates apparently affect lime
dosage required and can cause super-saturation of precipitating species. Typically, lime
dose required to reach a given pH is higher than predicted from softening equations.
Likewise, calcium and magnesium content of the product water is higher than would be
predicted for lime treated natural waters.
A correlation for estimating lime dose needed to bring a secondary effluent to a given
pH is presented in Figure 11-2. It is based on data from several sources (23-27). The
pH is reasonably well correlated with the ratio of lime dose to initial alkalinity. Other
factors which affect the result are ratio of magnesium to calcium, ionic strength, and
the phosphate level. The effects caused by these variables can be estimated, and computer
programs have been developed to make the necessary calculation (22) (28). However,
the accuracy of these calculations is provisional because of the uncertain effects of
contaminants in wastewater.
Figure 11-2 can be utilized to predict lime dose if an estimate of the desired pH level
can be made. The curve is entered at the ordinate, and the abscissa, the ratio of CaO
dose to alkalinity, is read from the graph. If alkalinity of the water is known, active
CaO dose can be calculated. Total quicklime dose can be calculated from active CaO
dose and the active CaO content of the quicklime.
11.3.2.3 Sludge Production
The principal reactions for estimating lime sludge production axe:
11-20
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FIGURE 11-1
EFFECT OF pH ON TOTAL RESIDUAL PHOSPHORUS
1.0
0)
E
o.i
' I
A POMONA
ODX BISHOP (BLUE PLAINS)
V BERG (LEBANON)
• NEVADA POWER CO.
< SOUTH TAHOE
D
LOG P=3.51 - 0.392(pH)
9.0
O
10.0
11.0
12.0
PH
11-21
-------
FIGURE 11-2
RATIO OF LIME DOSE TO ALKALINITY (BOTH MG/L)
FOR A PRESCRIBED FINAL pH
1 1 1
12
a:
LU
10
I 9
a
I 3
ALKALINITY
SYMBOL ALKALINITY REF LOCATION
(mg/l)
100-125
226
440-460
277
450
300
2 WASH., D.C.
4 S. TAHOE, CA
5 NINE SPRINGS, Wl
3 LEBANON, OH
6 LEBANON, OH
6 BATAVIA, OH
0.3 0.5 0.7 1.0 2.0
CaO/Alkalinity
3.0
5.0
11-22
-------
5Ca(OH)2 + 3HP04~- —- Ca5OH(PO4)3 + 6OH~ + 3H20 (1)
Mg++ + 2OH~ —•> Mg(OH)2 (2)
Ca(OH)2 + HCO3" —" CaCO3 + H2O + OH" (3)
The information required and the steps to take in estimating the sludge production are
listed below. Informational input needs are marked with an asterik.
TABLE 11-7
INFORMATION NEEDED TO CALCULATE LIME SLUDGE MASS
Information Needed
A. P*in> desired P removal*
(sometimes, BOD^, desired BOD
removal). Results of jar test
or experience*
B. Alkalinity of wastewater*, pH
level, Figure 11-2
C.
D.
, P removal, Equation 1
, from analysis or estimate*,
j, from analysis or
Figure 11-4, Equation 2
E. Ca^, from analysis or estimate*,
CaQUf, from analysis or
Figure 11-3,
Ca from lime dose (see item B),
Ca removal in HAP (see item C),
Equation 3
F. pH (see item A)
Determination
pH level
CaO dose
Ca5OH (PO4)3(abbrev.:HAP)
precipitated
Mg(OH)2 precipitated
CaCO3 precipitated
Verify that sufficient CO3
is present. If not,
estimates of Caou^ were in
error.
11-23
-------
G. Active CaO in lime* Inerts brought in with
quicklime.
The procedures for steps A and B have already been presented. Step C requires no
explanation. If estimates of Mgou^ and Caouf. are needed for steps D and E, available
data (24) can be used as shown in Figure 11-3 and Figure 11-4. They show the calcium
and magnesium levels obtained at the clarifier outlet in long-term evaluation of tertiary
lime treatment of secondary effluent. Magnesium (Figure 11-4) precipitation commences
at pH 10 and is essentially complete at pH 11. The percent in solution could be estimated
as declining from 100% at pH 10 to 0 at pH 11.
The calcium concentration (Figure 11-3) is seen to decline precipitously when pH is
increased to 9.5. Concentration then increases as pH is increased further. This behavior
is rational in light of knowledge of bicarbonate-carbonate equilibria. When pH is increased,
bicarbonate ion is converted to carbonate ion. When the solubility product of CaCOg
is exceeded, precipitation occurs and Ca concentration falls. As pH is increased further
by more Ca(OH)2 addition, eventually all of COj is removed by CaCO^ precipitation.
Calcium concentration then increases as calcium hydroxide dissolves to the limit of its
solubility.
Step E outlines the calculation of calcium carbonate precipitation. The calcium carbonate
deposited is calculated from a calcium balance. After considering the amount of calcium
removed in the hydroxyapatite (Ca5OH (PC^^), the CaCC^ is calculated from the net
loss of calcium.
A verification step is advisable to determine whether there is sufficient becarbonate to
account for the CaCOg produced. Bicarbonate can be determined from alkalinity. At
pH below about 8.3, all of alkalinity measured is due to bicarbonate. At pH below 7,
there is more potential carbonate available due to the presence of dissolved carbon dioxide,
which is not measured in the alkalinity test. Estimates of dissolved CC>2 can be made
by following procedures outlined in "Standard Methods" (29). If the bicarbonate is
insufficient to balance with lost calcium calculated from the material balance, calcium
in the final effluent was estimated at too low a value. There is little likelihood of this
problem at pH levels below about 10.8.
The entire procedure described, including the verification step and the totaling of all of
the components of the sludge, is shown in the illustration in Table 11-8. It is evident
from the total quantities listed at the end of the table that calcium carbonate makes
up about three-quarters of the mass of the sludge. For this reason, substantial errors
could occur in estimates of the other sludge quantities without introducing much error
into the final result.
-------
FIGURE 11-3
PRECIPITATION OF CALCIUM BY LIME ADDITION
100
O)
E
Zc*
w <
-------
FIGURE 11-4
PRECIPITATION OF MAGNESIUM BY LIME ADDITION
40
O)
E
O £
30
< u
U.
LU 20
10
I I
AVG INFLUENT .*v^
CONCENTRATION.^r-'-.
8
10
11
11-26
-------
TABLE 11-8
CALCULATION OF LIME SLUDGE MASS
Illustration of Procedure for Estimating Sludge Quantity
Secondary Effluent Characteristics
Phosphate Content 10 mg/1 as P
Alkalinity 300 mg/1 as CaCO3
Calcium 100 mg/1
Magnesium 20 mg/1
Treatment Conditions
Remove 85% of phosphate by raising pH to 10.5
Final Phosphorus
Content 1.5 mg/1 as P
Estimate of Ca
(Figure 11-3) 50 mg/1
Estimate of Mg
(Figure 11-4) 10 mg/1
Lime Dose
From Figure 11-1, 0.92 mg CaO/mg alkalinity at pH 10.5.
Dose = 0.92 x 300 = 276 mg/1 CaO
Assume quicklime is 90% active CaO
Total quicklime - 176/0.90 = 307 mg/1
Inerts with quicklime = (0.10)(307) = 30.7 mg/1
Sludge from Phosphate
(10 - 1.5)(502/93) = 45.8 mg/1 Ca5(OH)(PO4)3
Mg(OH)2 in Sludge
(20 - 10)(58.3/24.3) = 24.0 mg/1 Mg(OH)2
11-27
-------
Sludge from CaCOj Precipitation
Input Ca = 100 from sec. effluent + (276)(40/56) from lime dose
Output Ca = 45.8 x (5)(40)/(502) + 50 + Y
(removed with (remaining (precipitated
hydroxyapatite) in solution) as CaCO3)
Y = 100 + 197 - 18.3 - 50 = 129 mg/1
129 x (100/40) = 323 mg/1 CaCO3 precipitated
Total Sludge
Inerts in quicklime 30.7
Ca5OH(PO4)3 45.8
Mg(OH)2 24.0
CaCO3 323.0
Total = 433.5 mg/1
Verification that Sufficient CO3 is Present
300 mg/1 alkalinity (as CaCO3)
300 x (1/50) = 6.0 g-equivalents (or g-mols) HCO3
HCO3 + Ca"1"1" + OH" —•- CaCO3 + H2O
6.0 g-mols HCO3 can produce 6.0 g-mols CaCO3
6.0 x 100 = 600 mg/1 CaCO3
The CO3 available could produce as much as 600 mg/1 CaCO3. This
is substantially greater than the amount calculated above (323 mg/1)
from the calcium balance. Thus, there is ample CO3 present. If there
were insufficient CO3 (which can happen at ph of 11 and greater),
it would indicate that the estimate of Ca++ after treatment was too
low. The amount of CaCO3 indicated by the calculation from the
alkalinity would be correct.
11-28
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Additional Sludge from Recarbonation
Recarbonation to pH 9.5 will reduce Ca concentration to about
30 mg/1 (see Figure 11-3). For this example, the additional calcium
carbonate sludge produced is calculated as follows:
(50 - 30)(100/40) = 50 mg/1 CaCG>3
11.4 Dewatering at Specific Locations
11.4.1 Tertiary Treatment with Lime
11.4.1.1 South Tahoe
At South Tahoe, California, wastewater is treated in conventional primary, followed by
a waste activated sludge plant. Phosphate is removed by tertiary lime treatment to pH 11
with recarbonation. A comprehensive report describing the performance of the Tahoe
plant and sludge dewatering aspects has been published (25).
The tertiary lime sludge is dewatered and classified at Lake Tahoe by means of solid
bowl centrifuges. Sludge is thickened to 8 percent solids. It then flows at about 20 gpm
to a 75 HP Bird 24" x 60" solid-bowl concurrent-flow centrifuge. The centrifuge is
deliberately operated with high losses to the centrate in order to reject magnesium
hydroxide, hydroxyapatite and organic solids, leaving behind a high concentration of
calcium carbonate in the cake. The cake is calcined and is recovered. The centrate is
sent to a second centrifuge, polymer is added, and solids are removed as a 30% solids
cake.
The fact that a centrifuge tends to reject fine low-density solids may lead to fears that
fine solids will build up in the system. This has apparently not occured at South Tahoe.
11.4.1.2 Blue Plains Pilot Plant
At the Blue Plains pilot plant, a substantial amount of information was obtained on the
sludges produced by treating secondary effluent with lime. This information will be
published soon as a report in the EPA series (30).
Secondary effluent from the District of Columbia's Blue Plains Treatment Plant was treated
with lime to a pH range of 11 to 12 for phosphate removal. Instead of utilizing second-stage
recarbonation to remove excess Ca ions and recover additional calcium carbonate, sodium
carbonate was added to the single clarifier for this purpose. The sludge produced was
continuously thickened, vacuum filtered, calcined, slaked and recycled to the tertiary
clarifier to supply the bulk of the lime demand. Information on sludge thickening and
dewatering was obtained with and without recycle of reclaimed lime.
11-29
-------
Most of the data during recycle operations were obtained at 56 percent recovery; that
is, 56 percent of the active lime dose consisted of recycled lime. Recycle was
approximately 3 pounds of recycled solids to one pound of fresh lime. Lime dose
consisted of 300 mg/1 of active lime plus 240 mg/1 of inerts.
Buildup of inerts did not affect phosphorus removal but increased solids loading through
the sludge handling system. Solids content of sludge removed from the clarifier was 20 g/1
for no-recycle operation and 32 g/1 for 56 percent recycle.
The sludge was thickened in a 6-foot diameter vessel equipped with pickets. Sludge
thickened to 15-20% solids for recycle as well as no-cycle operations, despite solids loading
rates over 70 percent higher for the recycle case.
Information on filtration was obtained from filter leaf tests. Equations for form rate
for the recycle and no-recycle cases were developed. These equations, rearranged slightly
to give results in terms of yield and cycle time, are:
No Recycle: Y = (2.18) (f)°-5(Q °'84 (p) °'21/ (0) °'5
Recycle: Y = (3.28) (f) °-5 (C) °-80 (p) °-26 / (0) °-5
Where Y = yield, Ib/hr-sq ft
0 = cycle time (min), or reciprocal of rpm
C = Solids concentration, wt %
p = operating vacuum, in. Hg
f = fraction submergence
Calculating filter yield at the following conditions,
f = 0.25
C = 15% solids
p = 15" Hg vacuum
0=1 min. (i.e., 1 rpm)
11-30
-------
gives the following result:
No Recycle: Y = 18.7 Ib/hr-sq ft
Recycle: Y = 29.0 Ib/hr-sq ft
This study reports substantial cake cracking which reduced vacuum across the filter.
Nevertheless, solids contents from 30-40% were generally obtained. It was also noted
that, after a process upset which caused biological solids to carry over into the tertiary
treatment system, filter yield deteriorated drastically. However, filtration was possible
without chemical conditioning.
11.4.1.3 Butler, Indiana
At Butler, wastewater is treated in a 0.4 mgd conventional primary plus trickling filter
plant. Primary settling precedes two-stage biological filters. Conventional sludge is digested
and dried on beds. Phosphorus is removed by tertiary lime in a reactor-clarifier at pH 10.
Chemical dose is 150 mg/1 quicklime plus ferric chloride and polymer to aid settling.
Total phosphorus removal is about 80 percent. Sludge is not dewatered and is pumped
to a lagoon.
The Butler plant is somewhat unique in that the wastewater is quite hard (300 mg/1)
and its magnesium content is high (30 mg/I). Ordinarily, good phosphorus removal is
achieved in water of this type at a lower pH than in soft low-magnesium water. Required
phosphorus removal was only 80% and a pH of 10 was sufficient to get satisfactory
performance.
Sludges leaving the reactor-clarifier were about 8.4 percent solids. Overnight settling
increased this to 9.0 or 9.5 percent. Lagoon sludge, even after only a few days settling,
was about 17% suspended solids. Fresh sludge from the reactor-clarifier could be filtered
at 5.6 Ib/hr-sq ft at a cycle time of 2.9 min and 33% immersion. Higher yield could
be achieved at lower cycle times but this produced cakes less than one-eighth inch thick,
which caused problems with cake release.
Sludge was mostly mineral in nature. Percent volatile solids was under 10%. Magnesium
content of the sludge was low. At higher pH, more magnesium would have been
precipitated, which could have produced a poorer filtering sludge.
11.4.2 Lime to Primary Settler
11.4.2.1 Contra Costa
The Contra Costa Sanitary District, Walnut Creek, California, has carried out pilot and
plant scale evaluations of an integrated process for handling sludges produced by adding
11-31
-------
lime to primary wastewater treatment. This work has been reported (31) and is discussed
in Chapter 5.
The treatment process comprises lime flocculation with polymer or ferric chloride followed
by primary sedimentation, activated sludge under conditions such that nitrification is
achieved, denitrification with methane addition, and chlorination. Waste activated sludge
is recycled to the raw sewage to be removed from the primary clarifier along with the
chemical-primary sludge. Denitrification sludge is also returned to the incoming raw sewage,
but its amount is negligible.
In addition to investigations of process performance, studies were made to
evaluate (a) predictive methods for estimating sludge production, (b) the degree of
thickening achieved in primary clarification, (c) the utility of two-step procedure in which
sludge is centrifugally classified to recover calcium carbonate and the centrate is centrifuged
again to remove solids, and (d) anaerobic digestion of the first-stage centrate.
Sludge mass predicted from analysis of influent and effluent analyses checked satisfactorily
with measured sludge mass. Operation at pH 10.2 produced a clarifier sludge of about
4.0 percent solids. At pH 11, higher sludge solids could be attained. Centrifugal
classification achieved an 80% recovery of calcium carbonate in relatively dry cakes of
42-57% solids. Solids remaining in the centrate could be removed by polymer-assisted
second-stage centrifugation at recoveries as high as 90%. Results were not as good at
pH greater than 11. The centrates from the first centrifugation could be anaerobically
digested, although volatile solids reduction was lower than for conventional sludge
processing. In addition, precipitated magnesium redissolved in digester supernatant. Return
of supernatant to process would increase lime demand and magnesium would be
precipitated again. It is believed that handling of supernatant by separate treatment or
return to the secondary process would avoid that problem.
11.4.2.2 Eimco at Salt Lake City
Eimco Division of Envirotech has conducted a 100 gpm pilot plant investigation of a
wastewater treatment system in which chemicals are added in the primary step and soluble
organic material is removed by powdered activated carbon. A progress report on this
investigation has recently been published (14).
Comparisons of sludge mass calculated from clarifier input and output analyses with
measured mass showed reasonable agreement. When iron was added to the primary,
measured mass was 20% higher than the calculated value. However, this additional mass
appeared to include removal of soluble organics. When adding alum, calculated value
(including soluble organic material) was virtually the same as the measured value. For
lime, measured values averaged 20% higher than calculated values.
11-32
-------
Thickening and filtration data were incomplete for iron and alum. However, data with
lime were developed. After 24 hours thickening, sludges with pH values greater than
11.5 thickened to 2.0 to 9.8% solids, with median near 7%. For pH values less than
11.0, sludge thickened to 9.0 to 18.0% solids, with median near 14%.
Filtration yields were obtained for high and low pH operation. No polymer was needed.
For low pH sludge at the 14% thickened solids content, maximum filter yield was
6.0 Ib/hr-sq ft. For high pH sludge at the 7% thickened solids content, maximum filter
yield was 3.0 Ib/hr-sq ft.
It should be noted that Salt Lake City wastewater is high in magnesium. Operation at
pH greater than 11.5 precipitates much more magnesium than precipitated at pH less than
11.0. Since magnesium hydroxide is a flocculant precipitate, this could account for the
low solids content of thickened sludge from high pH operation.
11.4.2.3 Cleveland Westerly
Extensive pilot plant work was conducted at Cleveland's Westerly plant preliminary to
preparing a design for a physical-chemical treatment plant. A report describing this work
is in progress (32).
Most of the work was done with lime added prior to the primary to reach a pH of 10.5.
Sludge was dewatered using a Bird pilot plant centrifuge and an Eimco belt-type pilot
plant vacuum filter.
Sludge solids content leaving the clarifier was in the range of 6% to 14% solids. Thickening
prior to centrifugation was not needed. Solids recovery in the centrifuge was poor without
polymer. However, polymer doses of less than 2 Ib/ton gave recoveries of 95%. Solids
content of the sludge cake ranged from 20 to 28 percent.
The lime-primary sludge could be filtered without polymers but performance was improved
greatly by addition of small quantities of polymer. Filtration yields in excess of
10 Ib/hr-sq ft were recorded. Recoveries of solids were in the order of 95%. Cake solids
ranged from 20 to 35% with 28% a typical value.
11.4.2.4 Canadian Experience
The Province of Ontario's Ministry of the Environment has carried out a number of
plant-scale testing programs at sewage treatment plants. This work has been summarized
in several reports (33-35). Some of that information (34) is summarized here.
The use of lime in activated sludge plants increased sludge production from a normal
level of about 180 mg/1 to 460 mg/1 during addition of 200 mg/1 slaked lime. Sludge
solids content increased from 4.5% to 9%. In primary plants, sludge production increased
11-33
-------
from 80 mg/1 to 240 mg/1. Sludge solids content increased from 7% to 15%.
Some difficulty was encountered with anaerobic digestion because of erratic dosage with
high pH sludge. Subsequent studies have shown that holding the sludge longer in the
clarifier results in a fall in sludge pH. Digester operation is then normal except for high
solids content of the sludge.
Raw lime sludge was dewatered successfully in a continuous solids bowl conveyor-type
centrifuge at the Newmarket/East Gwillimbury plant (36). A detailed investigation showed
high recoveries could be achieved with polymer addition. At these high recoveries,
phosphorus removal from the centrate was also high (37).
With lime added to the primary of a conventional plant (primary plus activated sludge),
filter yield increased from 5.2 to 7.2 Ib/hr-sq ft, and conditioning costs (lime plusFeClj)
fell from $16/ton down to $1 I/ton. Cake solids content was 29%.
11.4.3 Alum and Iron Salts
11.4.3.1 Blue Plains
Alum was added to the 300 mgd wastewater treatment plant at Blue Plains, Washington,
D. C., for several months. At this plant, primary treatment is followed by high rate
activated sludge. Waste activated sludge is combined with sludge withdrawn from the
clarifier, and the mixed sludge is thickened and digested. The digested sludge was elutriated
in two stages using polymer to retain solids. The sludge was filtered using polymer and
ferric chloride for conditioning. The following information on alum addition tests is based
on discussions with plant operating personnel.
Alum was added to the aerator in gradually increasing dose until the average dose was
40 mg/1 of alum (9.1% Al). This developed a relatively high level of aluminum in the
sludge because the Blue Plains wastewater is weak.
Alum addition did not affect thickening. Sludge leaving the thickeners remained at 7.5%
solids. Digestion was likewise unaffected. Suspended solids leaving the digesters remained
at 3.5%. Elutriation could be carried out but a high polymer dose was required to retain
suspended solids. Chemical costs for elutriation increased from $3/ton before alum addition
to $7/ton during the test. Filter yields were about the same as with prior operation,
but solids content of the cake dropped from 23% to 20%. Polymer condition cost for
filtration, which had been $6/ton, rose to $7/ton.
11.4.3.2 Pomona
At Pomona, California, a physical-chemical plant has been in operation to produce a feed
stream for a long-term test of carbon columns. The wastewater has been treated in a
11-34
-------
primary clarifier with 22 mg aluminum per liter. A field test was conducted to determine
the dewatering characteristics of this sludge. That unpublished report (38) is summarized
here.
Gravity thickening tests showed that, without polymer, sludge settled slowly. Thickening
rate to produce a 4% solids sludge was quite low at 6.6 Ib/day-sq ft. With polymer
addition, that rate increased to 20 Ib/day-sq ft with 2 Ib/ton of anionic polymer and
to 70 Ib/day-sq ft with 6 Ib/ton. Filtration rates with 12% lime and polymer were less
than 2 Ib/hr-sq ft at the low cycle time of 2 min. However, the sludge employed was
only 2% solids. Rate would be expected to increase proportionately if thicker sludge
were available.
One of the reasons this field test was conducted was to verify other data indicating that
sludges produced when metals are added to the primary settle poorly and have lower
filter yields than conventional sludge. The Pomona field test corroborated these findings.
11.4.3.3 Grand Rapids
The addition of ferric chloride at various points in the process was studied at Grand Rapids,
Michigan. The Grand Rapids plant has a conventional primary followed by activated sludge.
The ferric chloride was added in three different modes: to the primary alone, split
between the primary and the aerator, and to the aerator alone. An investigation of sludge
dewatering was not part of this project. However, it was hoped that the filter yield
data routinely collected at this plant could be related back to the chemical dosing
conditions.
Monthly average filter yields, cake solids, and chemical costs were supplied by Grand Rapids
staff and are presented in Figure 11-5. Large fluctuations in all three filtration variables
are evident; however, they are not correlated with ferric chloride dosing. This finding
was predicted by Grand Rapids staff who observed that the basic nature of the sludge
can change from month to month and that necessary maintenance can cause step changes
in performance. Some of the interferences noted at Grand Rapids were:
a. Periodic dumping of lime sludge from a water plant into the
wastewater collection system.
b. Performance changes in digestion.
c. Acid cleaning of filter cloths.
d. Cleaning of interior piping of vacuum filters.
e. Replacement of vacuum system.
11-35
-------
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11-36
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f. Periodic switch-over from the incinerators, which had a lower
capacity than the vacuum filters and needed a dry cake, to landfill,
where maximum output was desired and a dry cake was not so
important.
11.4.3.4 Canadian Experience
Reference has been made in Section 11.4.2.4 to the work reported by Ontario's Ministry
of the Environment. Additional information from a summary paper (34) is included here.
When metal salts were utilized in activated sludge secondary plants, the dry solids sludge
production increased 5% to 25%. Sludge concentration dropped slightly to 3.5 to 5.0%
from the normal 4.0 to 5.5%.
In primary plants, sludge production increases substantially because of increased solids
capture. An increase of 100% was typical and 20% reduction in sludge solids concentration
was reported.
In all cases where metal salts were used to precipitate phosphorus, normal digestion was
observed. Problems which developed were the result of increased hydraulic and volatile
solids loading.
Filtering rates were approximately halved for a primary plant when alum was added. Lime
and ferric chloride conditioning costs increased from $3.10/ton to $9.50/ton. Solids
content of cake dropped from 31% down to 19%. For a secondary plant with alum
added to the primary, filtering rate fell from 5.2 to 4.8 Ib/hr-sq ft, and conditioning
chemical cost increased slightly. Cake solids was typically 15.9% with alum addition.
11..4.3.5 Chapel Hill
Chapel Hill has a 5 MGD trickling filter plant which is conveniently split into two parallel
treatment plants, each with its own primary clarifier, high rate trickling filter, and final
clarifier. During an experimental program, the wastewater from one train was dosed with
alum for phosphorus removal, just before it entered the final clarifier. A field trip was
arranged to this unique facility to determine the impact of alum addition on the wastewater
sludge. Unpublished results (39) are summarized here.
Data obtained indicated sludge from the secondary clarifier following the alum-dosed train
was essentially unfilterable, even with polymer addition. Since this was a high-rate filter
with substantial recirculation before the clarifier, the sludge was very low in organic matter.
The control secondary clarifier sludge showed reasonable filter yields. Secondary sludges
from both trains were recirculated to their respective primary clarifiers and removed as
mixed primary-secondary sludges. Filter leaf tests showed that there was no substantial
difference in the filter yields from the two trains.
11-37
-------
Data supplied by University of North Carolina staff who operate the plant indicated that
4.75 Ib extra sludge was generated per Ib of Al added. They also indicated that difficulty
was encountered during alum trials in their two-stage digester: the solids content of the
supernatant leaving the second digester was unusually high.
11.5 References
1. "Process Design Manual for Sludge Handling and Disposal," Technology Transfer,
U. S. Environmental Protection Agency (1974).
2. Smith, R., Eilers, R., "Wastewater Treatment Plant Cost Estimating Program,"
NTIS PB 219472 (April, 1971).
3. Yeaple, D. S., Barnes, D. A., and DiGiano, F. A., JBF Corp., "A Computer
Model for Evaluating Local Phosphorus Removal Strategies," Environmental
Protection Agency Contract 68-01-0758, Draft Report (January 3, 1973).
4. Mogelnicki, S., Los Angeles Technology Transfer Session (November, 1972).
5. Teletzke, G. H., "Sludge Dewatering by Vacuum Filtration," Rocky Mountain
Sew. & Ind. Wastes Assoc. Annual Meeting, Colorado Springs, CO
(October 24-26, 1960).
6. Adrian, D. D., Smith, J. E., Jr., "Dewatering Physical-Chemical Sludges,"
Vanderbilt Symp., "Applications of New Concepts of Physical-Chemical
Wastewater Treatment," in press (September 18-22, 1972).
7. Isgard, E., "Chemical Methods in Present Sewage Purification Techniques," 7th
Effluent and Water Treatment Exhib. and Conv., London (June 25, 1971).
8. Mulbarger, M. C., Shifflett, D. G., "Combined Biological and Chemical Treatment
for Phosphorus Removal," Chem. Engr. Prog. Symp. Ser. 67, No. 107, 107-116
(1970).
9. Farrell, J. B., "Alum-Phosphate Sludges: Dewaterability and Aluminum
Recovery by Lime Treatment," In-House Report, Adv. Waste Treatment Research
Lab., NERC, Cincinnati, OH, U. S. Environmental Protection Agency (June,
1973).
10. Smith, R., "Estimating the Rate of Sludge Production in the Activated Sludge
Process," Memo (October 10, 1972).
11-38
-------
11. City of Austin and University of Texas, "Design Guides for Biological Wastewater
Treatment Processes," U. S. Environmental Protection Agency Water Pollution
Control Research Series 11010 ESQ (August, 1971).
12. Kreissl, J. F., Westrick, J. J., "Municipal Waste Treatment by Physical-Chemical
Methods," Vanderbilt Symp., "Applications of New Concepts of
Physical-Chemical Wastewater Treatment," (September 18-22, 1972).
13. "Process Design Manual for Phosphorus Removal," U. S. Environmental
Protection Agency (October, 1971).
14. Burns, D. E., Shell, G. L., "Physical-Chemical Treatment of Municipal Wastewater
Using Powdered Carbon," Final Report, U. S. Environmental Protection Agency
Project No. 17020 EFB (1972).
15. Friedman, L. D., Weber, W. J., Bloom, R., and Hopkins, C. B., "Improving
Granular Carbon Treatment," Water Pollution Control Research Series 17020
GDN (1971).
16. Farrell, J. B., unpublished data.
17. Westrick, J. J., Cohen, J. M., "Comparative Effects of Chemical Pretreatment
on Carbon Adsorption," 47th Annual Conf., Water Pollution Control Federation,
Atlanta, GA (October, 1972).
18. Schaffer, R. B., Van Hall, C. E., McDermott, G. N., Barth, D., Stenger, V. A.,
Sevesta, S. J., and Griggs, S. H., "Application of a Carbon Analyzer in Waste
Treatment," JWPCF, 37:11, p. 1545-56.
19. Bunch, R. L., Barth, E. F., and Ettinger, M. B., "Organic Materials in Secondary
Effluents," JWPCF, 33:2, p. 122-26 (February, 1961).
20. Brown, J. C., "Alum Treatment of High-Rate Trickling Filter Effluent at Chapel
Hill, North Carolina," Durham, NC (April, 1973).
21. Derrington, R. E., Stevens, D. H., and Laughlin, J. E., "Enhancing Trickling
Filter Plant Performance by Chemical Precipitation," U. S. Environmental
Protection Agency, Grant #11010 EGL, (July, 1973).
22. Seiden, L., Patel, K., "Mathematical Model of Tertiary Treatment by Lime
Addition," Report No. TWRC-14, FWPCA (September, 1969).
23. O'Farrell, T. P., Bishop, D. F., and Bennett, S. M., "Advanced Waste Treatment
at Washington, D. C.," "Water-1969," Chem. Engr. Prog. Symp. Ser. 65, No. 97,
11-39
-------
p. 251-58.
24. Berg, E. L., Brunner, C. A., and Williams, R. T., "Single-Stage Lime Clarification
of Secondary Effluent," Water & Wastes Engr., (March, 1960).
25. South Tahoe Public Utility District, "Advanced Waste Water Treatment as
Practiced at South Tahoe," Water Pollution Control Research Series 17010 ELQ,
p. 89 (August, 1971).
26. Malhotra, S. K., Lee, G. F., and Rohlich, G. A., "Nutrient Removal from
Secondary Effluent by Alum Flocculation and Lime Precipitation," Internatl,
Jour. Air-Water Poll, 8, p. 487-500 (1964).
27. Mulbarger, M. C., Grossman, E. Ill, Dean, R. B., and Grant, O. L., "Lime
Clarification, Recovery, Reuse, and Sludge Dewatering Characteristics,"
Jour WPCF, 41:12, p. 2070-85 (1969).
28. Smith, R., "Calculation of Ionic Equilibria by Means of the Digital Computer,"
Advanced Waste Treatment Research Lab., Cincinnati, OH (November, 1966).
29. "Standard Methods for the Examination of Water and Wastewater,"
13th Edition, APHA, AWWA, WPCF, Washington, D. C. (1971).
30. Bennett, S. M., Bishop, D. F., "Solids Handling and Reuse of Lime Sludge,"
Environmental Protection Agency Contract #14-12-818, 11010 FYM.
31. Parker, D. S., Zadick, F. J., and Train, K. E., "Sludge Processing for Combined
Physical-Biological Sludges," Environmental Protection Agency, R2-73-250 (July,
1973).
32. Zurn Engineers and Battelle/NW, "Westerly Advanced Waste Treatment
Facility: Process Development and Engineering Design," Draft Version
(September 16, 1971).
33. Van Fleet, G. L., Barr, J. R., and Harris, A. J., "Treatment and Disposal of
Chemical Phosphate Sludges in Ontario," Water Pollution Control Federation
Meeting, Atlanta, GA (October, 1972).
34. Boyko, B. I., Rupke, J. W. G., "Design Considerations in the Implementation
of Ontario's Phosphorus Removal Program," Phosphorus Removal Design
Seminar, Toronto, Ontario, Ministry of the Environment (May 28-29, 1973).
11-40
-------
35. Knight, C. H., Mondoux, R. G., and Hambley, B., "Thickening and Dewatering
Sludges Produced in Phosphorus Removal," Phosphorus Removal Design Seminar,
Toronto, Ontario, Ministry of the Environment (May 28-29, 1973).
36. Black, S. A., "Lime Treatment for Phosphorus Removal at the Newmarket/East
Gwillimbury WPCF, an Interim Report," Ontario Ministry of the Environment,
R.P.W. 2032 (May, 1972).
37. Smith, A. G., "Centrifuge Dewatering of Lime Treated Sewage Sludge," Ontario
Ministry of the Environment, R.P.W. 2030 (May, 1972).
38. Hathaway, S. W., Unpublished Report for EPA/NERC (August 27, 1973).
39. Hathaway, S. W., Unpublished Report for EPA/NERC (September, 1972).
11-41
-------
APPENDIX A
EVALUATING PHOSPHORUS REMOVAL STRATEGIES
A computer model for evaluating a number of strategies for removing phosphorus in
wastewater has been developed for EPA.1 This model reports to the user the total cost of
a selected strategy for removing phosphorus. 22 treatment schemes can be selected and
evaluated depending on local conditions. These schemes are:
1. Alum & Polymer Addition to Primary Settler
2. Ferric Chloride & Polymer Addition to Primary Settler
3. Ferrous Chloride & Lime Addition to Primary Settler
4. Lime Addition to Primary Settler
5. Alum & Polymer Addition to Flocculation Basin
6. Ferric Chloride & Polymer Addition to Flocculation Basin
7. Ferrous Chloride & Lime Addition to Flocculation Basin
8. Lime Addition to Flocculation Basin
9. Alum Addition to Aeration Basin
10. Ferric Chloride Addition to Aeration Basin
11. Sodium Aluminate Addition to Aeration Basin
12. Alum Addition to Aeration Basin & Multi-Media Filtration
13. Ferric Chloride Addition to Aeration Basin & Multi-Media Filtration
14. Sodium Aluminate Addition to Aeration Basin & Multi-Media Filtration
15. Alum Addition After Trickling Filter
16. Ferric Chloride Addition After Trickling Filter
1 7. Alum Addition After Trickling Filter & Multi-Media Filtration
18. Ferric Chloride Addition After Trickling Filter & Multi-Media Filtration
19. Alum Addition to Flocculation Basin After Conventional Secondary Treatment
20. Ferric Chloride Addition to Flocculation After Conventional Secondary Treatment
21. Lime (1-Stage) Addition to Flocculation After Conventional Secondary Treatment
22. Lime (2-Stage) ^ Addition to Flocculation Basin After Conventional Secondary
Treatment
Four levels of effluent phosphorus concentration (total unfiltered phosphorus) were
assumed to be achievable depending upon the type of treatment provided. These are:
2.0 mg/1 (Option 1) Chemical addition to the primary clarifier. (Treatment
schemes 1-4)
(Option 2) Chemical addition to a flocculation basin prior to the primary
clarifier. (Treatment schemes 5-8)
(Option 3) Chemical addition to the aeration basin. (Treatment schemes
9-11)
1 "A computer Model for Evaluating Community Phosphorous Removal Strategies," prepared for U.S. EPA
by JBF Scientific Corp., Burlington, Mass., Report No. 400/9-73/001 (Oct. 1973).
A-l
-------
(Option 4) Chemical addition after the trickling filter. (Treatment
schemes 15-16)
0.5 mg/1 (Option 5) Chemical addition to the aeration basin plus multi-media
filtration. (Treatment schemes 12-14)
or
Chemical addition after the trickling filter plus multi-media filtration.
(Treatment schemes 17-18)
(Option 6) Addition of lime to a flocculation basin following conven-
tional secondary treatment. (Treatment scheme 21)
0.3 mg/1 (Option 7) Addition of alum or ferric chloride to a flocculation basin
following conventional secondary treatment—one stage. (Treatment
schemes 19-20)
0.1 mg/1 (Option 6) Addition of lime to a flocculation basin following conven-
tional secondary treatment—two stages. (Treatment scheme 22)
The model has been designed specifically for planning purposes and can be used by
municipalities, planning agencies and government policy organizations. In order to insure its
usefulness, EPA plans to periodically update the model as additional performance data on
phosphorus removal processes becomes available.
The phosphorus removal model (REMOVE) is designed to provide an estimate of the cost to
a particular community of achieving a designated level of phosphorus removal. The program
begins with the entry of data describing the community in terms of population and future
trends, the treatment plant in terms of present size and excess capacity, and the influent
wastewater characteristics. The data can be entered interactively through a subroutine
(SYSTEM).
After input data have been entered, the user has the opportunity to investigate various
control strategies for achieving a low phosphorus concentration in the wastewater effluent.
Typically, the effects on cost of reducing phosphorus in detergent products, of increasing or
decreasing the amount of government financing available, or of using various treatment
options can be investigated.
It is important to note that the model has been developed as a simulation only and as such
does not achieve a mathematical optimization.
From the description of the community, treatment plant, wastewater characteristics, and
desired strategy the model will design a treatment plant incorporating the present plant as
described by the user and will add those processes or additional equipment necessary to
achieve the desired effluent level. The program converts the sizing information into
construction costs, operating man-hours, maintenance man-hours, and total material and
supply costs for each process.
A-2
-------
The following pages illustrate the type of print-out obtained when using the computer
model. The particular cases illustrated here are a 5 mgd trickling filter plant and a 20 mgd
activated sludge treatment plant.
5 MGD TRICKLING FILTER PLANT
TREATMENT COSTS IN CENTS/1000 GAL
P
GOVFF
RESTRICT
SCHNO
1
2
3
k
5
6
7
8
9
10
11
12
13
16
17
18
21
7.7
8.2
2
.8
6.7
7.3
2
.P
yes
0.5
13
10
17
13
Ik
11
18
Ik
12
10
15
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11.2
10.5
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10.2
9.1
11.5
7.7
11.2
NOTES:
1. P = Phosphorus effluent level in ms/1,
2. GOVFF = Fraction of construction cost financed
by federal and state governments.
3. RESTRICT = Restrictions on phosphate decereents
k. SCHNO = Treatment scheme number.
5. = Lowest treatrrent cost.
A-3
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PERIODIC OPERATING COSTS
PROCESS NAME
RAW WSTE PNiP
PRE TRETMENT
PRIM SETTLER
PRMYSLGE PMP
TRKLNG FILTR
SECY SETTLER
RECIRC PMPNG
CHLORNE FEED
CHLORNE BASN
GRAV THICKNR
SLGE HLD TNK
VACUUM FILTR
ALUM FEEDING
FECL FEEDING
LIME FEEDING
TOTALS =
OPERATING MAN-HOURS
(MAN-HOURS/TP)
AVE. MAX. MIN.
MAIMTAiriAHCE MAN-HOURS
O-'AN-HOURS/TP)
AVE. MAX. MIN.
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u,:v!T COST DATA
(CENTS/ijiC GALLONS)
PROCESS NAME AMORTIZATION
LAF'IR
MATERIAL « SUPPLY TOTAL
RAW WSTE PKP
PRE TRETMENT
PRIM SETTLER
PRMYSLGE PMP
TRKLNG FILTR
SECY SETTLER
RECIRC PMPNG
CHLORNE FEED
CHLORNE BASN
GRAV THICKNR
SLGE HLD TNK
VACUUM FILTR
ALUM FEEDING
FcCL FEEuING
LIME FEEDING
TOTALS =
TIME SINK Af
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A-7
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p
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SCMNO
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TRtATMLNT COSTS IN
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yes
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7.1
NOTES:
1. P = Phosphorus effluent level in mr;/1 .
2. GOVFF = Fraction of construction cost financed
by federal and state governments.
3. RESTRICT = Restrictions on phosphate duter ;ents
k. SCHNO = Treatment scherre number.
D. = Lowest treatment cost.
A-8
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PERIODIC OPERATING COSTS
PROCESS NAME OPERATING MAN-HOURS MAINTAINANCE MAN-HOURS
(KAN-HOURS/TP) (VAN-HOURS/TP)
AVE. MAX. MIN. AVE. MAX. MIN.
RAW WbTE PMP 0. 0. 0. 0. C. 0.
PRE TRETMENT 0. 0. 0. 0. C. 0.
PRIM SETTLER 0. 0. 0. 0. 0. 0.
PRMYSLGE PMP 0. 0. 0. 0. G. 0.
AERTN BASIN 0. 0. 0. 0. 0. 0.
MECH AERATN 0. 0. 0. 0. 0. 0.
SECY SETTLER 0. 0. 0. 0. 0. 0.
RECIRC PKPNG 0. 0. 0. 0. 0. 0.
CHLORNE FEED 0. 0. 0. 0. 0. 0.
CHLORNE BASN 0. 0. 0. 0. 0. 0.
GRAV THICKNR kl. l»l. 1*1. 23. 23. 25.
SLGE HLD INK 0. 0. 0. T. 0. 0.
VACUUM FILTR 3013. 3013. 3013. 3f9. 3f,9. 359.
FECL FEEDING 0. 0. 0. 285U. 2S8U. 2££i*.
LIME FEEDING 0. 0. 0. 2003. 2003. 2009.
TOTALS = 3055. 3055. 3055. 52.U. 5 2 P I*. 52£1*.
A-10
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0..'.
UNIT COST DAT-\
(CEMS/100G GALLONS)
PROCESS NAME AMORTIZATION LAP.OP, MATERIAL i SUPPLY TOTAL
C.I
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RAVj w'STE PMP
PRE TRETMENT
PRIM SETTLER
PR.'-'.YSLGE PMP
AERTN BASIN
KECH AERATN
SECY SETTLER
RECI3C PM.PNG
CHLORNE FEED
CHLORNE 8ASN
GRAV THICKNR
SLGE HLD TNK
VACUUM. FILTR
FECL FEEDING
LIME FEEDING
TOTALS •
TIME SINK
0
0
0
0
A
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0
0
0
0
0
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0
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.52
SOL
AMOUNT
PERIOD USED
1
2
':>
k
S
6
7
£
9
10
11
12
13
11*
15
16
17
18
19
20
TOTALS «
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
1
in s
SOL
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
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.21
.18
.12
.52
COS
TS
ins
(TON'S) (
5
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
2553
1079
.9
.9
.9
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.9
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7
7
7
7
7
7
7
7
7
7
7
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7
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7
7
7
7
5
COST
?/TP)
C3RU9.
638U9.
638
638
638
653
638
63S
658
63S
638
63?,
638
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638
638
638
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12
769
85.
36
36
36
36
36
36
36
36
56
56
36
36
36
36
36
36
3C
36
36
36
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0.05
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5.31
UNIT COSTS FOR SOLID WASTE DISPOSAL (CENTS/1000 GAL.)
TOTAL UNIT COST (CENTS/1000 GAL) = 6.18
= D.R7
A-12
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Multiply
Inches
Feet
Square Feet
Cubic Feet
Pounds
Gallons
Gallons/Minute
Feet/Second
APPENDIX B
METRIC CONVERSION CHART
By To Get
2.54
0.3048
0.0929
0.0283
0.454
3.79
5.458
0.305
Centimeters
Meters
Square Meters
Cubic Meters
Kilograms
Liters
Cubic Meters/Day
Meters/Second
Environmental
Lakes National fi
GLHPO
B-l
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