PROCESS DESIGN MANUAL FOR PHOSPHORUS REMOVAL
For
U. S. ENVIRONMENTAL PROTECTION AGENCY
Technology Transfer
By
BLACK & VEATCH, CONSULTING ENGINEERS
P. O. Box 840S
Kansas City, Missouri 64114
Program #17010 GNP
Contract #14-12-936
October, 1971
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The mention of trade names or commercial products in this manual is
for illustration purposes, and does not constitute endorsement or
recommendation for use by the Environmental Protection Agency.
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CONTENTS
Chapter Page
INTRODUC 1ON
1.1 Purpose I - 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 — Chapter 2 2 - 5
3 THEORY OF PHOSPHORUS REMOVAL
BY CHEMICAL PRECIPiTATiON
3.1 Forms and Measurement of Phosphorus 3 -
3.2 Chemistry of Removal by Precipitation 3 - 2
3.3 References — Chapter 3 3 - 6
4 PHOSPHORUS REMOVAL BY
MINERAL ADDITION BEFORE THE PRIMARY SETTLER
4.1 General Considerations 4 -
4.2 Phosphorus Removal Data for Alum Addition 4 - 4
4.3 Phosphorus Removal Data for Iron Addition 4 - 6
4.4 Overall Performance of Treatment Plants
with Mineral Addition to the Primary Settler 4- 11
4.5 Sludge Handling Requirements 4 - 13
4.6 Dosage Selection and Correlation 4 - 16
4.7 Plant Designs for Phosphorus Removal with Iron Addition 4- 16
4.8 Costs 4 - 22
4.9 References — Chapter 4 4 - 24
5 PHOSPHORUS REMOVAL BY
LIME ADDITION BEFORE THE PRIMARY SETTLER
5.1 Description of Process 5 - 1
5.2 Process Performance 5 - 1
5.3 Example of a Plant Design 5 - 3
5.4 Capital and Operating Costs 5 -6
5.5 References — Chapter 5 5 7
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Chapter Page
6 PHOSHPORUS REMOVAL IN
TRICKLING FILTERS BY MINERAL ADDITION
6.1 Pre-Design Decisions 6 - 1
6.2 Process Options 6 - I
6.3 Performance Data Using Aluminum Salts 6 - 2
6.4 Performance Data Using Iron Salts 6 - 5
6.5 Choice of Chemical Addition 6 - 6
6.6 Nature and Role of Chemicals Involved 6 - 7
6.7 Dosage Selection and Control 6 - 8
6.8 Iron Leakage 6 - 9
6.9 Sampling and Analysis 6 - 9
6.10 References — Chapter 6 6 - 14
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 - 23
7.5 References — Chapter 5 7 - 32
8 PHOSPHORUS REMOVAL BY
LIME TREATMENT OF SECONDARY EFFLUENT
8.1 Description of Process 8 - I
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 — Chapter 6 8 - 21
9 PHOSPHORUS REMOVAL BY
MINERAL ADDITION TO SECONDARY EFFLUENT
9.1 Description of Process 9 - I
9.2 Summary of Design Information 9 - I
9.3 Laboratory and Pilot Studies 9- 2
9.4 References — Chapter 9 9 - 7
10 STORAGE AND FEEDING OF CHEMICALS
10.1 Aluminum Compounds 10 - 1
10.2 Iron Compounds 10 - 15
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Chapter Page
10 STORAGE AND FEEDING OF CHEMICALS — (Continued)
10.3 Lime 10-23
10.4 Other Compounds for pH Adjustment 10 - 32
10.5 Carbon Dioxide 10-41
10.6 Polymers 10- 46
10.7 RapidMixing 10-49
10.8 References — Chapter 10 10- 52
10.9 Supplemental Bibliography 10- 52
11 CHECKLISTS FOR PLANS AND SPECIFICATIONS
11.1 Introduction 11-1
11.2 Phosphorus Removal by
Mineral Addition Before the Primary Settler 11 - I
11.3 Phosphorus Removal by
Lime Addition Before the Primary Settler 11 - 3
11.4 Phosphorus Removal in
Trickling Filters by Mineral Addition 11 - 5
11.5 Phosphorus Removal in
Activated Sludge Plants by Mineral Addition 11 - 5
11.6 Phosphorus Removal by
Lime Treatment of Secondary Effluent 11 - 6
11.7 Phosphorus Removal by
Mineral Addition to Secondary Effluent ii - 8
APPENDIX A -
ACKNOWLEDGMENT AC - 1
vii
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FOREWORD
The formation of the Environmental Protection Agency marks a new era of environ-
mental 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 municipal 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 manualS is one of four now available through the sponsorship of the Environ-
mental Protection Agency to describe recent technological advances and new informa-
tion in the following subject areas:
Carbon Adsorption
Phosphorus Removal
Upgrading Existing Plants
Suspended Solids Removal
These manuals are the first edition copies and will be updated as warranted by the
advancing state of the art to include new data as it becomes available, and to refine
design criteria as additional full-scale operational information is generated.
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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.
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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.! 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 defici nt 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 lb per capita per year (2). The mean annual excretion is estimated to be
about 1.2 lb per capita. The mean annual contribution of phosphorus from synthetic
detergents with phosphate builders is estimated to be about 2.3 lb 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 lb 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 5,000 x 3.5 = 17,500 lb/year, or
48 lb/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/l(asP)
The average total phosphorus concentration in domestic raw wastewater is found to be
about 10 mg/I 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
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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 (Ii). 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/I as P, with many established at 1.0 mg/i. 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 (lb/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/I
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Receiving Water - (continued)
Maximum allowable phosphorus concentration 0.05 mg/i
Low flow 140 ft 3 fsec (90 nigd)
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-I-4.8)
X = 0.8 mg/i 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 — Chapter 2
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-
erice 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 & md Wastes,
29:2, p 157 (1957).
9. Lawton, G. W., et al., “Spray Irrigation of Dairy Wastes”, Sew. & md. Wastes,
31:8, p 923 (1959).
10. Flynn, J. M., and Andres, B., “Launderette Waste Treatment Processes”, JWPCF,
35:6, p 783 (1963).
11. Wymore, A. H., and White, J. E., “Treatment of Slaughterhouse Waste Using Anaer-
obic and Aerated Lagoons”, Wa :. & 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 P0 4 3 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 UP0 4 2 . With the present practice of reporting phosphorus as
P, one does not usually have to be concerned with the actual form. lt 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.
Dunng 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 hoc. 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 (I), 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 A nalysis of Water
and Wastes (2)
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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
form, PO 4 3 , 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 floc 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:
A1 3 ÷ P0 4 3 AIPO 4
The above equation indicates that the mole ratio for AI:PO is I : 1. Inasmuch as the
mole ratio of P:P0 4 is also 1:1, the mole ratio for Al:P is 1:1 or —p- 1 when both
aluminum and phosphorus are expressed in terms of gram-moles or lb-moles. On a
weight, rather than a mole basis, this means that 27 lb of Al will react with 95 lb of
P0 4 to form 122 lb of AIPO 4 . Since each 95 lb (1 lb-mole) of P0 4 contains 31 lb
(1 lb-mole) of P, the weight relationship between Al and P is 27 lb of Al to 3 1 lb 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 Al 2 (S0 4 ) 3 14H 2 0 (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 A1 2 O 3 or
about 9.1% expressed as Al. Its reaction with P0 4 3 may be written as follows:
Al 2 (SO 4 ) 3 14H 2 0 + 2P0 4 3 —’2A1P0 4 + 35042 + 14H 2 0
The sulfate remains in solution as S042. The above reaction indicates that 1 lb-mole
of alum (594 lb) will react with 2 lb-moles (190 Ib) of P0 4 3 containing 62 lb
phosphorus to form 2 lb-moles (244 Ib) of AIPO 4 . 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 Al:P mole ratio as follows:
Mole ratio Al:P = 1:1
Therefore weight ratio Al:P = 27:31 = 0.87:1
Alum contains 9.1% Al 0 7
Therefore, alum required per lb of P O. l 9.6 lb
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Stumm and Morgan (3), state that the solubility of AIPO 4 is p1-I dependent and varies
as follows:
Approximate
pH Solubility (mg/I )
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 p 1-I 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
At 2 (S0 4 ) 3 • l4H 2 O + 6HCOj —÷2Al(OH) 3 + 6C0 2 + l4H 2 O + 3SO42
The following ratios of alum (9. l%Al) to phosphorus are believed reasonably repre-
sentative for alum treatment of municipal wastewater (4).
Al: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: ! 16:1
95% 23:1 2.0:1 22.1
To achieve 85% P removal from a wastewater containing 11 mg/I of P the alum dosage
needed would be (16)(ll) 176 mg/I or 1470 lb/b 6 gal.
Sodium aluminate will also serve as a source of aluminum for the precipitation of
phosphorus. The chemical formula for sodium aluminate is Na 2 A1 2 O 4 or NaAlO 2 .
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One commercial form is the granular trihydrate, which may be written Na 2 0 . A 1 2 0 3 3H O
and which contains about 46% A1 2 0 3 or 24% Al.
The scidium aluminate - phosphate reaction may be represented as follows:
Na 2 O Al 2 O 3 + 2PO 4 3 + 4H 2 0 > 2AlPO + 2NaOH + 60Ff
In contrast to alum, which reduces pH, a nse in pH may be expected upon addition
of sodium aluminate to wastewater.
The above reaction indicates an Al:P mole ratio of 1: 1, an Al: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 (Fe 2 ) and ferric (Fe 3 ) ions can be used in the precipitation of
phosphorus With Fe 3 + a reaction can be written similar to that shown earlier for
precipitation of aluminum phosphate. A 1: 1 mole ratio of Fe.P0 4 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 Fe 2 + the situation is more complicated and
not fully understood. Ferrous phosphate can be formed. The mole ratio of Fe:PO 4
would be 3.2. Experimental results indicate, however, that when Fe 2 + is used, the mole
ratio of Fe:P will be essentially the same as when Fe 3 + 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 Fe 3 +
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
Fe 2 + 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 2 + 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 HP0 4 2 .
3 HPO 4 2 + 5Ca 2 + 40H > Ca 5 (OH)(P0 4 ) 3 + 3 H 2 0
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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 — Chapter 3
1. Standard Methods for the Examination of Water and Wastewater, 13th Edition,
American Public Health Assoc., New York (1971).
2. FWPCI4 Methods for Chemical Analysis of Water and Wastes, Federal Water
Pollution Control Administration, U.S. Department of the Interior (Nov 1 1969).
3. Stumm, W and Morgan, J J., Aquatic Chemistry . p 521 Wiley-lnterscience,
John Wiley & Sons, Inc., New York (1970)
4. A]hed 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 mclude 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 I 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 carned 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.
4-I
-------
WEST PRIMARY ——
TO ________
SECONDARY r
CLARIFIER _______
AERATION 1 A FI WE ELL CHEMICAL
I SLUDGE
I PRE-AERATION J
TAN K
POLYMER
__________ __________ ADDITION
I’ _
RAW _____________________________
— PARSIIALL FLUME F
WASTEWATER __________ __________
ALUM OR FERRIC CHLORIDE
ADDI lION
FIGURE ‘ 4-I LEBANON, OHIO PRIMARY TREATMENT AND CHEMICAL FEEDING SCHEMATIC
-------
BENTON HARBOR INTERCEPTOR
FeC 3 ADDITION POINT,
I NTERCEPTOR MANHOLE
ST. JOSEPH
INTERCEPTOR
a
COMM I NUTOR
RAW SEWAGE
PUMPS
VENTURI
TO PRIMARY
DOW A-23
ADDITION POINT
FIGURE ‘4-2 SCHEMATIC OF PRETREATMENT OF PLANT FLOW AT BENTON HARBOR -
ST. JOSEPH WASTEWATER DISPOSAL PLANT
-------
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
tnckling 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 Al: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/I by 80%.
Raw wastewater at Two Rivers required an average Al:P weight ratio of 0.93:1 (molar
ratio of 1.07:1) to reduce the average influent phosphorus of 20.1 mg/I by 85%. These
test data indicate that as phosphorus concentration increases, the Al: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/I, and the AI: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/I, as A1 2 (S0 4 ) 3 . 14H 2 0, and anionic polymer
dosages of 0.25 and 0.5 mg/I should be evaluated. Only the Parshall flume,
preaeration basin, and primary clarifiers of the I .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/I was found to
be satisfactory for improving hoc formation and settling. An alum dose of 250 mg/I,
AI:P weight ratio of 3.4:1 (molar ratio of 3.9:1), removed 93% of the phosphorus in
this high alkalinity (400 mg/i as CaCO 3 ) wastewater. Effluent phosphorus
concentration did not exceed 0.7 mg/I and averaged 0.5 mg/I. The alum dose of 150
mg/I, Al: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/I, Al:P weight ratio of 3.9 I (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/i dosages
studied.
Little data are available for sludges resulting from alum addition to primary settlers. At
Lebanon, Ohio, the primary sludge produced by alum addition received lime
treatment before final disposal as landfill or on farm land (1 2). Before lime treatment,
the sludge was similar in character to biological sludge and had a solids concentration
of from 2 to 3%. It had very poor settling and thickening characteristics and was
malodorous.
Lime treatment consisted of mixing the sludge with a 25% lime slurry until a pH of
11 .5 was reached. The average lime requirement was 440 lb of lime as Ca(OH) 2 per
ton of dry solids. A contact time of 30 minutes was maintained before the sludge was
allowed to flow to sand beds for drying. Seventy per cent of the total volume of
sludge applied to the drying beds was removed as drainage in the first 3 days, and the
sludge cake could be lifted off the beds in 3 weeks. Elimination of odors and
4-5
-------
complete pathogen kill was accomplished with lime treatment. Lime costs at $ 18/ton
for Ca(OJ-1) 2 , arid an application rate of 440 lb/ton of dry solids, resulted in a
chemical cost for sludge handling of $3.96/ton of dry sludge solids.
4.3 Phosphorus Removal Data for Iron Addition
A number of studies have been conducted using the addition of iron to raw
wastewater for phosphorus removal. These investigations have ranged from laboratory
jar testing to full plant-scale work and have included both ferrous and ferric salts as
the source of iron. Average values for doses, phosphorus removals, and references to
the studies are listed in Table 4-1.
These studies include wastewaters with phosphorus concentrations of 6 to 9 mg/I, so a
range of phosphorus removals would be expected. In addition, there are many other
significant variations in wastewater character from location to location. The results of
the studies allow bracketing the phosphorus removals obtained and enable drawing
conclusions based on special features of individual studies.
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/I, as
Fe. The Escanaba study showed that base and polymer addition could halve the iron
dose required to obtain the same phosphorus removal achieved with iron alone. The
Lebanon study showed that doubling the iron dose (from 45 to 90 mg/I) with the
same polymer addition raised phosphorus removal from 68 to 91%. At Wyoming,
Michi an, the same phosphorus removal was obtained with equal doses of either Fe 2 +
or Fe +. Polymer doses in all of the iron studies ranged from 0.26 to 0.7 mg/I.
Iron carry-over occurred at Mentor, where FeCI 2 addition alone resulted in an effluent
Fe concentration of 42.5 mg/I. Lime addition reduced the iron leakage to an effluent
Fe concentration of 11.6 mg/I. Polymers were ineffective in reducing iron leakage.
Increased suspended solids and HOD 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 HOD removal is seen in the Grayling, Michigan study.
Removal of HOD 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
hours at Benton Harbor. At Mentor, air at a tate of 42.5 to 85 ft 3 /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
4-6
-------
Table 4-1
STUDIES OF IRON ADDITION BEFORE THE PRIMARY SETFLER FOR PHOSPHORUS REMOVAL
Fe Dose (mg/I)
and/or Base Polymer % P Removal
Location Type of Study Iron Form Fe:P (by weight ) (mg/I) (mg/I) in Primary
Grayling,
Michigan (13) plant scale FeCI 2 15-25 30-50 0.3 60-80
(NaOH) (Dow A23,
anionic)
Lake Odessa,
Michigan (14) plant scale FeCI 2 15-25 30-50 0.3 75* .93*
(NaOH) (Dow A23,
anionic)
Bay City,
Michigan (15) lab FeCl 2 10, 1.8:1 30 0.5 80
(CaO) (Dow A23,
anionic)
Escanaba, Fe 3 2.2:1 0 0 80
Michigan (16) lab Fe 3 1.1: 1 40 0.35 80
(CaO) (Dow A23,
anionic)
Wyoming,
Michigan lab Fe 2 10 30 0.7 80
(17,18,19) (CaO) (Dow A23,
anionic)
Fe 3 10 0.4 80
(Dow A23,
anionic)
-------
Table 4-1 (Continued)
Fe Dose (mg/I)
and/or Base Polymer % P Removal
Location Type of Study Iron Form Fe:P (by weight ) (mg/I) (mg/I) in Primary
Benton Harbor- lab FeC! 3 20 0.3 optimum
St. Joseph, (Dow A23,
Michigan (20) anionic)
5.5 mgd FeC! 3 21 0.26 89.7
(activated sludge (Dow A23,
return to primary) anionic)
5.5 mgd FeCI 3 21 0.26 65.3
(no sludge return (Dow A23, 90.9*
to primary) anionic)
3
Cleveland, 35 mgd-Imhoff Fe + 20 0.5 78***
Ohio (21) tanks as (Dow A23,
primaries anionic)
Warren,
Michigan (22) 36 mgd FeCI 3 15 80*
Lebanon, 1.01 mgd FeC! 3 45, 1.76:1 0.5 68
Ohio (11) (Atlas Altasep
2A2, anionic)
0.86 mgd FeC I 3 90, 3.46:1 0.5 9!
(Atlas Altasep
2A2, anionic)
-------
Table 4-1 (Continued)
Fe Dose (mg/i)
and/or Base Polymer % P Removal
Location Type of Study Iron Form Fe:P (by weight ) (mg/I) (mg/I) in Primary
Mentor, lab FeCI 2 60 80 >80
Ohio (waste pickle) (CaO)
(23,24) 4 mgd FeCI 2 3.1 : 1 CaO:Fe, 80
(waste pickle) by weight
was 1.9:1, to
adjust pH to
7.5-8.0
*Secondary treatment provided. Removal value is for entire plant.
**Operated at 5.5 mgd. Design capacities were 8 mgd for primary and 4 mgd for secondary, “Kraus Process”, treatment.
***% reduction in effluent P.
-------
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. At Bay City, Michigan, the optimum
time intervals in the chemical addition sequence were found to be 10 seconds between the
iron and lime addition followed by 5 minutes mixing before the addition of Dow
Purifloc A23 polymer.
Information related to iron sludges and sludge handling was obtained at Lake Odessa
and Grayling, Michigan, and Cleveland, Lebanon and Mentor, Ohio. The sludge
resulting from the addition of iron for phosphorus removal can be handled by
anaerobic digestion, vacuum filtration, incineration, a combination of these three, or
lime treatment. Each of the first three methods has been used during plant-scale
studies of iron addition. The last method was used on a research study during a
plant-scale addition of iron.
At Lake Odessa, anaerobic sludge digestion is practiced and digester supernatant was
returned to the primary unit during the test period. The total phosphorus
concentration in the supernatant decreased during the study from an initial
concentration of 100 to 200 mg/i to 23 mg/I during latter stages of the study
period, and it was concluded that phosphorus release was not occurring during
digestion. A crystalline phosphate precipitate found in the raw sludge from Grayling
and the digested sludge from Lake Odessa was identified by X-ray diffraction
techniques as vivianite, Fe 3 (P0 4 ) 2 .8H 2 0. Analyses by X-ray diffraction also showed
that the two sludges contained 20 to 24% iron.
The chemical sludge resulting from iron addition at Mentor, Ohio (23, 24), was
digested anaerobically and dewatered on vacuum filters. The volume of sludge
produced during the iron and lime treatment was never more than double that of
periods without chemical treatment. Iron-lime primary sludge averaged 8.5 to 10.0%
solids in November and December, 1969. The anaerobic digesters operated with normal
gas production and pH. The larger volumes of sludge shortened digester residence time
but stabilization did occur. Typical concentrations of iron and phosphorus in the
various sludges are given below in Table 4-2.
Table 4-2
TYPICAL IRON AND PHOSPHORUS ANALYSIS
DURING SLUDGE HANDLING AT MENTOR, OHIO
Primary Sludge Analysis Digester
Raw Digested Supernatant
(mg/l) (mg/I) (mg/I)
Total Iron 3750 1020 3950
Soluble Iron 105 26 trace
Total Phosphorus (as P) 1140 393 1240
Soluble Phosphorus 81.5 32.7 6.6
4 - 10
-------
At Mentor the iron-lime digested sludge had good dewatering characteristics. The
digested sludge with 1 0% solids was dewatered on a vacuum filter to an average of
20% solids.
At Cleveland, Ohio (21) anaerobic digestion, vacuum filtration, and incineration were
utilized to dispose of the 28 tons/day of solids produced by Fed 3 addition to the
primary units. Polymer conditioning of both raw and digested sludge resulted in filter
yield rates of up to 10 lb/ft 2 /hr. No problems were encountered in incinerating the
cake.
Lime treatment of ferric chloride sludge at Lebanon, Ohio (12) was designed on the
basis of 8,700 gpd of sludge. Lime addition to the sludge to a pH of II .5 required
244 lb of lime, as Ca(OH) 2 , per ton of dry slu4ge solids. After 30 minutes of contact
at this pH, the sludge was pumped to drying beds. Samples of limed and unlimed
sludge were analyzed, and corresponding sludge characteristics are given in Table 4-3.
The analyses show that lime treatment increased sludge solids and filter yields but had
little or no effect on filter cake moisture content.
Table 4-3
SLUDGE FROM FERRIC CHLORIDE TREATMENT
OF RAW SEWAGE AT LEBANON, OHIO
Before Lime After Lime
Sludge Solids, % 1.9 2.4
Volatile Portion, % 70 57
Dissolved Solids, % 0.17 0.17
Vacuum Filter Yield*
90 mg/I FeCI 3 dose 1.06 1.57
45 mg/I FeCI 3 dose 1.57 2.40
Filter Cake Moisture Content**
90mg/I FeCl 3 dose 4.10 4.28
45 mg/I FeCl 3 dose 3.75 3.75
*lb of sludge solids/ft 2 /hour
**lb of water/lb of sludge solids
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 4-4. Phosphorus removal without
chemical precipitation in the primary clarifiers ranges from only 10 to 20% through a
4-Il
-------
Table 4-4
EFFECTIVENESS OF PRIMARY AND SECONDARY TREATMENT WITH
AND WITHOUT MINERAL ADDITION FOR PHOSPHORUS REMOVAL
Phosphorus Removal Suspended Solids Removal BOD Removal
(%) (%) (%)
Without With Without With Without With
Primary Treatment 5-10 70-90 40-70 60-75 25-40 40-50
Secondary Treatment
Trickling Filter 10-20 80-95 70-92 85-95 80-90 80-95
Activated Sludge 10-20 80-95 85-95 85-95 85-95 85-95
-------
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. Thc 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
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 (25).
4.5 Sludge Handling Requirements
Application of phosphorus removal during primary treatment will result in increased
amounts of primary sludge because of the greater capture of suspended solids and the
precipitated metal phosphate and hydroxide solids. Consequently, the amount of
secondary sludge will generally decrease, and the ratio of primary to secondary sludge
will increase.
Sludge conditioning characteristics will also be changed. Instead of requiring cationic
polymers or ferric chloride and lime for conditioning, sludges resulting from mineral
addition to the primary settler will probably best be conditioned with an anionic
polymer. Filter yields should remain about the same as those for sludges from
conventional treatment without mineral addition. Addition of polymer to the raw
wastewater may reduce the polymer demand for conditioning the sludge.
4 - 13
-------
Anaerobic digestion will be unaffected by the addition of chemically precipitated
sludge. The higher digester, loadings resulting from additional sludge production will not
normally be detrimental to operation unless an organic overloading condition exists.
Release of soluble phosphorus from the sludge during digestion is considered to be
minimal. In situations requiring exceptionally high degrees of total phosphorus removal
(>90%), it may be desirable to treat sidestreams such as digester supernatant and
filtrate separately with metal salts before recycle to the head of the plant.
An example of estimation of the quantities of sludge and its distribution in a hypothetical
activated sludge plant are given in Table 4-5. Relative concentrations of suspended solids,
BOD, and total phosphorus are based on a hypothetical wastewater with process
efficiencies in the range of those given in Table 4-4.
Table 4-5
ESTIMATION OF SLUDGE PRODUCTION
Assumed Conditions Conventional Chemical
Suspended solids, mg/I (% removed)
Raw 300 300
Primary 150 (50) 75(75)
Secondary 30 (90) 15(95)
Soluble BOD, mg/I (% removed)
Raw 150 150
Primary 150(0) lS0( 0)
Secondary 30 (80) 10 (93)
Total phosphorus, mg P/I (% removed)
Raw 10 10
Primary 9 (10) 2 (80)
Secondary 8 (20) 1 (90)
Chemical Dose 25 mg Fe/I = 73 mg FeCI 3 /1
Chemical Precipitation
Fe 3 + P0 4 3 ) FePO
16.2 mg/I 9.0 mg/I P 44 mg/I
Fe 3 + 30H ) Fe(OH) 3
8.8 mg/I 8.1 mg/I 17 mg/I
73 mg FeCl 3 /l gives 44 mg FePO 4 /I + 17 mg Fe(OH) 3 /l
= 61 mg precipitated Fe solids/I
1 lb FeCI 3 gives 0.60 lb FePO 4 + 0.23 lb Fe(OH) 3
= 0.83 lb precipitated Fe solids
4 - 14
-------
Table 4-5 (Continued)
Sludge Balance Conventional Chemical
Initial suspended solids captured, lb/lU 6 gal.
Primary 1250 1875
Secondary 1000 500
Overall 2250 2375
BOO converted solids, lb/b 6 gal.
Primary 0 0
Secondary jQQ. 583
Overall 500 583
Precipitated iron solids, lb/i o6 gal.
Primary 0 448
Secondary _-.—Q
Overall 0 448
Total solids, lb/b 6 gal.
Primary 1250 2323
Secondary 1500 1083
Overall 2750 3406
Weight ratio of primary sludge to secondary sludge 0.83 2.15
% Change. chemical over conventional
Primary sludge + 86
Secondary sludge - 28
Total sludge + 24
The relative amounts of FePO 4 and Fe(OH) 3 are estimated from simplified stoichiometry.
Secondary or biological sludge quantities have been estimated on the basis that 0 5 lb
of solids is produced per lb of soluble BOD applied.
In comparing a conventional secondary treatment plant with the same plant having
chemical treatment in the primary for phosphorus removal, the addition of chemical
treatment results in better solids capture, improved BOO conversion, and greatly
increased phosphorus removal. The calculated ratio of primary to secondary sludge is
more than double that of conventional treatment, and in actual operation would result
in improved sludge conditioning The addition of chemical treatment also reduces the
load to the secondary, which should result in some improvements in actual operation.
Chemical treatment increases the total sludge mass only 24%, of which about 16% is
due to formation of the insoluble iron-phosphate-hydroxide sludge.
4 - 15
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4.6 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 I 0 for general
information on various chemicals. Material presented here will be limited to comments
and/or data on dosage selection, dosage correlation, and rnfluent 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/I to determine the required chemical dose in mg/I.
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.
4.7 Plant Designs for Phosphorus Removal
with iron Addition
The following three designs for proposed expansions of wastewater treatment facilities,
at Escanaba, Bay City, and Wyoming, Michigan, include iron addition for phosphorus
removal. All three were prepared by Black & Veatch, Consulting Engineers, and are
designed for iron addition to the primary basin. Escanaba and Wyoming will use ferric
chloride while Bay City will use ferrous chloride.
4.7.1 Escanaba, Michigan
A complete-mix activated sludge plant at Escanaba, Michigan is being constructed to
replace a combination trickling filter-conventional activated sludge plant. Provisions
have been made in the design to meet the State’s requirement of 80% phosphorus
removal. Effluent from the plant will flow to Little Bay DeNoc of Lake Michigan.
4- 16
-------
1.0
3
0.l
a
=
C ,)
a
=
LU
-j
-J
a
C,)
a
I-
C .)
0.01
a
LU
a
3
C -)
a
=
C ,,
a
=
LU
-j
-J
a
C,)
0 0.5 1.0 1.5 2.0 2.5 3.0 3.5 ‘hO 4.5
METAL TO INITIAL SOLUBLE PHOSPHORUS RATIO, WEIGHT BASIS
FIGURE ‘1-3 SOLUBLE PHOSPHORUS REMOVAL BY FERRIC CHLORIDE ADDITION
-------
+ ‘L96
MEAN 3. 58
—— -cr= 2.20
TIME, DAYS
FIGURE )4-14 DIURNAL VARIATION OF ORTHO PHOSPHORUS
8
7
6
AVERAGES FOR 5—DAY WEEK INFLUENT TO PLANT
FLOW = 5.15 mUd
BOD 95 mQ/i
SUSPENDED SOLIDS 182 m f I
U)
a
U)
a
0.
C
I-
C
3
2
MON I TUE 2 WED 3 JIIU ‘4 FRI 5 SAT 6 SUN 7
4 - 18
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Phosphorus removal in the new plant will be accomplished by precipitation with FeCI 3
in the primary basin, using polymer to enhance settleability of the floc. Ferric chloride
will normally be fed to the wastewater in a new rapid mix basin which will have a 3.4
minute detention time at the average design flow of 2.2 mgd. Alternate iron addition
points will be located at the raw wastewater pump station and the Parshall flume.
Polymer will normally be added toward the effluent end of an air flocculation basin,
having a 65 minute detention at design flow. Polymer may also be added at any of
the points of iron addition. Phosphorus precipitation will occur in the four primary
basins, each having a detention time of 3 hours at the design flow.
The storage facilities for FeCl 3 are designed to provide a 30 day supply at a feed rate
of 27 mg/I Fe 3 to a plant flow of 2.2 mgd. These facilities consist of two 4,000 gal.
rubber lined steel tanks, 8 ft in diameter and II ft high. It is anticipated that FeCI 3
will be stored and fed at the shipping strength, 35 to 45% solution. However, a
metered water supply will be connected to each storage tank to allow dilution, if
necessary. Two metering pumps will be furnished and piped to take suction from
either storage tank. Either pump will be capable of delivering Fed 3 over a range of
from 1.6 to 40 gal./hour. This will allow a feed rate range of from 10 mg/I (as Fe) of
45% solution at a minimum plant flow of 1 .0 mgd to 30 mg/I (as Fe) of 35% solution
at a peak hourly plant flow of 5.0 mgd. The storage tank and metering feed pump
area will be isolated to trap any leakage.
Bagged polymer storage, solution equipment, and metering pumps will be located near
the FeCl 3 facilities. Polymer solution will be made up in pre-selected concentrations by
an automatic volumetric batching system. A metering pump will deliver polymer from
a holding tank to any of three application points. This pump is sized for chemical
dosages ranging from 0.20 mg/I at 1.0 mgd to 0.30 mg/I at 5.0 mgd. Due to the high
viscosity, even of very dilute solutions, polymer (Dow A23) will be fed by a positive
displacement rotary gear pump. The feed range for a 0.5% solution will be 5 to 125
gal./hour.
Sludge from the primary basins will be pumped to the existing anaerobic digesters,
having a total capacity of 77,700 ft 3 . Digested sludge will be pumped to drying beds
with a total area of 37,090 ft 2 . Waste activated sludge will be digested aerobically.
4.7.2 Bay City, Michigan
The existing wastewater plant at Bay City provides only primary treatment before
discharge of the wastewater to Lake Huron through the Saginaw River. The State of
Michigan is requiring 80% phosphorus removal and improved organic removal. Iron
addition will be used for phosphorus removal and trickling filters will be provided for
organic removal. The expanded facilities are now under construction.
4 - 19
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The phosphorus removal process design is based on studies conducted by the Dow
Chemical Company. These studies indicated that the most efficient method of
removing phosphorus would consist of addition of FeCl 2 , a base (NaOH), and an
anionic polymer (Dow A23) to the raw wastewater. Although Dow’s recommendations
for Bay City’s raw wastewater indicated ferrous chloride to be the best iron salt,
ferrous sulfate or ferric chloride may also be used as alternatives.
Ferrous chloride will be furnished as a liquor containing a minimum concentration of
22% FeCl 2 by weight. The liquor will be shipped in 4,000 to 5,000 gal. tank trucks.
Two 20,450 gal. storage tanks will provide a minimum of 2 weeks supply of liquid
FeCI 2 . Feeding of ferrous chloride will be automatically paced with plant flow at
present, however, instrumentation is specified such that future control of the iron feed
rate in proportion to phosphate concentration is possible if a continuously monitoring
phosphate analyzer is added. Positive displacement metering pumps will be provided to
feed the liquid FeCI 2 into the wastewater at the upstream end of the Parshall flume at
the plant inlet.
Sodium hydroxide will be obtained as a 50% caustic soda solution in 3,000 to 4,000
gal. tank trucks. Two 6,450 gal. storage tanks will provide over a month’s supply of
50% NaOH, depending on the plant flow and alkalinity of the wastewater. Because
20% caustic solution has a much lower temperature of crystallization, two I ,200 gal.
day tanks equipped with mixers will be provided for diluting the 50% solution to 20%
before it is pumped from a heated chemical building to the treatment plant. A
centrifugal pump will transfer 50% solution to the day tanks when the storage tank
levels are low. Positive displacement metering pumps will feed NaOH to the wastewater
in proportion to plant flow at a rapid mixer at the downstream end of the Parshall
flume. The time of flow between the point of addition of FeCI 2 and NaOH is
approximately 10 seconds.
The chemically dosed wastewater will be divided and will flow into two new air
agitated mixing basins downstream from the Parshall flume. Each basin will be 1 7.5 ft-
wide by 38.7 ft long and will have a sidewater depth of 12.6 ft. At the maximum
plant flow rate of 18 mgd, approximately 10 minutes detention will be provided. At
the average plant flow rate of 1 2 mgd and with one basin out of service, there will be
approximately 7.5 minutes detention. Because some material may settle in these basins,
screw conveyors will be installed in each to remove such material to a sump. Settled
material will be pumped to the downstream side of these basins to be re-mixed with
the plant flow. Stop plates and drains will allow each basin to be removed rapidly
from service for maintenance.
The polymer facilities are designed for use of Dow Purifloc A23. Two 1,500 gal. solution
tanks, each equipped with a mixer, will be included. Two gear-type pumps will be used
to feed polymer in proportion to plant flow A third standby pump will also be provided.
Polymer will be added at rapid mixers downstream from the air agitated basins. The
wastewater will then flow into flocculation basins that are gently stirred by air for I to 5
minutes and then into settling basins. Three centrifugal blowers will be provided, one
4 - 20
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will supply air for agitation of the two mixing basins, one will supply air for agitation of
the flocculation basins, and the third will be a standby. A small air line will be run to the
ferrous chloride storage tanks to allow intermittent agitation. All chemical storage tanks,
day tanks, and pump feeders are located in a heated building.
It is anticipated that sludge from the expanded plant can be handled by existing
equipment. At present, sludge is stored in holding tanks until a sufficient amount has
accumulated for operation of the filters and incinerator. The sludge is conditioned
with polymer prior to filtration on two 150 ft 2 vacuum filters. Filter cake is then
incinerated in a 1 4.25 ft diam., six hearth incinerator.
4.7.3 Wyoming, Michigan
An expansion of the Wyoming, Michigan wastewater treatment plant is being designed
for 80% phosphorus removal and increased organic removal. The plant will include new
chemical mixing and flocculation basins, two primary settling basins, a chemical feed
and storage building, secondary aeration basins (trickling filters are to be used for
roughing), final settling basins, and sludge handling facilities. Phosphorus will be
removed by addition of ferric chloride to the new chemical mixing basins, followed by
polymer addition, flocculation, and sedimentation in the expanded primary facilities.
Provision will be made to feed ferric chloride and polymer to the final clarifiers for
increased phosphorus or solids removal. All chemical tanks and feed equipment will
also be designed to allow the use of liquid alum, ferrous chloride, ferrous sulfate, and
sodium hydroxide.
Plant components associated with phosphorus removal are to be designed on the
following basis. The chemical mix and flocculation basins will have a combined
detention time of 20 minutes at the average design flow of 19 mgd. At design flow,
the two existing and two new primary basins will have overflow rates of 604 gpd/ft 2 ,
weir rates of 15,100 gpd/ft, and detention times of 2.07 hr.
Capacity of the sludge handling facilities will be more than doubled. The present units
consist of two holding tanks with mixers, two 250 ft 2 vacuum filters, and one 6
hearth incinerator rated at 1,500 lb dry solids/hr. The expanded sludge processing
facilities have been designed on the following basis:
Suspended Solids
Plant load, lb/day 63,460
Overall removal (95% assumed), lb/day 60,290
Biochemical Oxygen Demand
Plant load, lb/day 48,260
Removed by aeration basins, lb/day 7,850
Waste activated sludge, Ib/day* 3,925
Total 64,215
4 - 21
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Solids added by lime and ferric chloride conditioning (15%) 9,635
Total solids loading, lb/day 73,890
Total solids loading, lb/hour 3,080
*Waste activated sludge production based on 0.5 lb solids generated per lb
BOD removed.
Additional equipment which will be provided to handle this load includes two air
flotation thickeners, two sludge holding tanks, two vacuum filters, and one incinerator.
The two air flotation thickeners will each have an effective area of 1,170 ft 2 . The
sludge will be conditioned with lime or ferric chloride prior to flotation thickening at
a rate of 20 lb/ft 2 fday. The two holding tanks will allow 4 days storage capacity for
the additional sludge. Each of the two new 500 ft 2 vacuum filters will have a capacity
of 4.2 lb/ft 2 /hr, and one will be on a standby basis. The six hearth incinerator will
handle 2,100 lb of dry solids per hr. Lagoons will be provided for disposal of the
ash.
4.8 Costs
All cost estimates for the removal of phosphorus by chemical precipitation-flocculation
during primary treatment are dependent to some extent on raw wastewater flow rates.
The cost per 1 ,000 gal. for chemicals may be less in larger installations because of
large volume purchasing at lower unit prices. Total capital costs for tankage, chemical
feed pumps, and instrumentation will be less for smaller plants. Placed on a unit
volume of waste treated basis, however, capital costs for small plants will be higher
than for large plants
4.8.1 Chemical Costs
The chemical costs for phosphorus removal in the primary settler have been computed
for both liquid alum and liquid FeCl 3 . The respective unit costs of these chemicals
were estimated at $60/ton for alum (on a 17.1% A1 2 0 3 basis) and $90/ton, as FeCI 3 ,
for liquid FeCl 3 . These figures represent average delivered prices.
As a basis for estimating costs in terms of / I ,000 gal., an influent P concentration of
10 mg/I was assumed. For removal of 80% of the total phosphorus in pnmary
treatment, equal molar ratios of Al:P and Fe:P of 1.5: 1 were assumed. Polymer costs
were included, based on a dosage rate of 0.5 mg/I and a chemical cost of $1.50/lb.
The chemical costs of phosphorus removal treatment are estimated to be 3.6 4/1,000
gal. for tiquid FeCl 3 and 4.2 4/1,000 gal. for liquid alum, respectively. As mentioned
earlier, these costs may vary depending on the size of treatment plant and cost of
freight.
The chemical cost for phosphorus removal at plants where waste pickle liquor is
available is more difficult to predict accurately because of varying iron concentrations,
larger storage requirements, freight charges, and the additional alkalinity requirement.
4 - 22
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4.8.2 Capital Costs
Equipment costs for phosphorus removal with aluminum and iron are small relative to
the chemical costs. Costs for facilities will be presented in Chapter 10 for single point
addition of several different chemicals. For convenience, however, capital costs of a
typical primary addition process using liquid ferric chloride and polymer are given
below in Table 4-7.
Table 4-7
STORAGE AND FEEDING CAPITAL COSTS FOR
LIQUID FERRIC CHLORIDE AND POLYMER ADDITION
BEFORE THE PRIMARY
Plant Size Capital Cost
(mgd) (Total $) (c / I ,000 gal.)
20,810 0.446
10 71,740 0.154
100 697,000 0.149
*Based on 25-year amortization at 6%.
The capital cost expressed on a /l,000 gal. basis allows comparison with the chemical
costs given earlier and illustrates that the capital costs are not very significant. Capital
costs for alum addition are the same or less than those for liquid ferric chloride
addition, depending on plant size. Further information on costs arid on the design
basis adopted for determining capital costs is given in Chapter 10.
4 - 23
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4.9 References — Chapter 4
I. 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 Waler 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.
Pmg. 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, i. 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 - 24
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13. Green, 0., Eyer, F., and Pierce, D., “Studies on Removal of Phosphates and
Related Removal of Suspended Matter and Biochemical Oxygen Demand at
Grayhng, Michigan, March — September 1967”, Division of Engineering, Michigan
Department of Health, Lansing, Michigan, and The Dow Chemical Company,
Midland, Michigan (1967).
14. ______________, “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 Department
of Public Health, Lansing, Michigan, and The Dow Chemical Company, Midland,
Michigan (1967).
15. Dow Chemical Company, Private Communications (1968).
16. Black & Veatch, Unpublished Data ((1970).
17. 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).
18. “Report on Wastewater Treatment Plant Improvements, Wyoming, Michigan”, Black
& Veatch of Michigan, Consulting Engineers, Kansas City, Missouri (1970).
19. Stonebrook, W. J., Dykhuizen, V., Beegh ly, J. H., and Pawlak, T. J., “Phosphorus
Removal at a Trickling Filter Plant Wyoming, Michigan”, Unpublished (Undated).
20. Johnson, E. L., Beeghty, J. H., and Wukasch, R. F., “Phosphorus Removal with
Iron and Po lyelectro lytes”, Public Works, 100:11, p 66 (1969).
21. Cherry, A. L., and Schuessler, R. G., “Private Company Improves Municipal Waste
Facility”, Wat and Wastes Eng, 8:3, p 32 (1971).
22. Boggia, C., and Herriman, G. L., “Pilot Plant Operation at Warren, Michigan”,
Proceedings of the 43rd Annual Conference of the Michigan Pollution (‘on trot
Association (1968).
23. 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, Prepubtication
Copy (1971)
24. Progress Reports, Mentor, Lake County, Ohio, FWPCA Grant WRPD 172.01-68,
May 15, 1968 through May 31, 1969 and September I, 1969 through July 31,
1970
25. Voshel, D., and Sak, 3 C., “Effect of Primary Effluent Suspended Solids and
BOD on Activated Sludge Pioauccion”, JWP(’F, 40.5, Parr 2, p R203 (1968)
4 - 25
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Chapter 5
PHOSPHORUS REMOVAL BY LIME ADDITION
BEFORE THE PRIMARY SETTLER
5.1 Description of Process
Addition of lime to raw wastewater for the purpose of phosphorus precipitation can
be applied to conventional plants using the activated sludge process. Because the pH of
the primary effluent ahead of the activated sludge process should not exceed a value
of 9.5 to 10.0, phosphorus removal by this method is generally limited to efficiencies
of about 80%. Higher pH values can result in biological upset. Greater phosphorus
removal, if necessary, may be accomplished by aluminum or iron addition to the
aerator or final settling basin. One fortunate aspect of the lime clarification-activated
sludge process is that the high pH of the primary effluent is reduced in the aeration
tank by the CO 2 produced by biological metabolism.
The removal of phosphorus in primary treatment is accompanied by increased BOD
and suspended solids removal and thereby reduces the organic load on the secondary
treatment facilities. Reduced organic load benefits overloaded plants and also aids in
establishing nitrification. Some phosphorus removal is accomplished by biological
treatment and this method takes advantage of that characteristic to remove some of
the remaining phosphorus that would be more difficult to remove by chemical
treatment. Use of lime rather than an iron or aluminum salt does not add large
amounts of other ions such as sulfate or chloride to the treated water. Other possible
benefits include improved oil, grease, and scum removal and less corrosion in the
primary sludge handling system.
One consideration in selecting the lime clarification-activated sludge process is the
ability to use existing primary settling basins for phosphorus removal. The initial
capital expenditure for phosphorus removal, therefore, may be comparatively small.
There is a need for feeding lime, mixing of the lime with the water, and flocculation.
Lime feeding equipment is discussed in Chapter 10. Equipment requirements for
mixing and flocculation will vary depending on the configuration of the existing plant.
in some cases, mixers and flocculators may have to be provided. In others, satisfactory
flocculation will occur with no additional equipment.
Sludge disposal may become more complex and lime sludge may not be compatible
with existing disposal processes such as anaerobic digestion. Furthermore, recovery of
lime by recalcimng which is discussed in Chapter 8 may not be practical with this
process because of the low CaCO 3 content of the sludge.
5.2 Process Performance
The concept of using lime for phosphorus removal in primary treatment preceding
activated sludge was advanced a number of years ago. Since then it has been
investigated at laboratory scale and full plant scale.
5-1
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Albertson and Sherwood (I) established with laboratory data that lime treatment of
raw wastewater is an efficient means of removing 80% or more of the phosphorus and
60 to 70% of the BOD. Discharge of the pH 9.5 or above water to a complete-mix
activated sludge system eliminated the need for pH correction because of CO 2
production by the biological activity in the aerator.
Schmid and McKinney (2) investigated lime precipitation of phosphate from wastewater
on a laboratory scale. Their studies also included precipitation of various phosphate
compounds from distilled water solutions by the addition of lime and consideration of
the dewatering characteristics of the sludge formed by phosphate precipitation. These
investigators concluded that both orthophosphate and polyphosphates, normally present
in wastewater, can be precipitated readily with lime. They also noted that the complex
phosphates present in wastewater seriously nhibit calcium carbonate precipitation. In
view of this, they concluded that sludge recalcination for lime recovery would not be
practicable. For the wastewater employed in the test work, the addition of 150 mg/I
of lime (calcium hydroxide) resulted generally in a pH of 9.5, total phosphorus
removal before biological treatment of 60%, and suspended solids removal of 90%. The
authors concluded that sludge produced by primary treatment, where lime is employed,
may be about twice that obtained normally. Overall sludge production, including waste
activated, would be about 1.5 times normal. The process can be controlled by
operating at a constant pH. Operation of a complete-mix activated sludge system is not
hindered, provided proper control is exercised. Microbial CO 2 production in
the activated sludge unit is sufficient to maintain the pH near neutral. A mixture of
lime-precipitated primary sludge and waste activated sludge will dewater and filter well
with small additions of anionic polymers.
Following extensive laboratory jar test and bench-scale studies, Black and Lewandowski (3)
conducted a plant-scale investigation of lime treatment for phosphorus removal at the
180,000-gal. (imperial) per day Richmond Hill, Ontario activated sludge plant. Lime
was added before the primary clarifier on a continuous basis for approximately 3
months. In this case, no special mixing or flocculating equipment was provided. The
raw wastewater had an alkalinity of 381 mg/l as CaCO 3 , a total hardness of 401 mg/I
as CaCO 3 , and an average total phosphorus content of 10.6 mg/I as P. Normal plant
operation produced a total phosphorus removal of 39% with BOD and suspended solids
removals in the primary unit of 21% and 37%, respectively. A lime dosage of 175
mg/I was sufficient to produce a pH of 9.3. During lime addition, BOD and suspended
solids removals in the primary clarifier were increased to 72 and 78%, respectively, and
total phosphorus removal averaged 83%. Phosphorus removal over the total plant was
92% resulting in a total phosphorus concentration in the secondary effluent of 0.9
mg/I. Throughout the study, CaCO 3 accumulated on the effluent weirs and outlet
troughs of the primary clarifier but was easily removed with a water hose. There was
no appreciable buildup on submerged structures within the primary unit. Lime
treatment Shad no detrimental effects on the secondary process. The high pH of the
primary effluent was promptly reduced by the CO 2 generated in the aeration tank.
Sludge equilibrium was reached and maintained throughout the study. As a result of
improved primary performance, the organic loading to the secondary plant was reduced,
5-2
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resulting in its behavior as an extended aeration process with no secondary sludge
being wasted. If sludge wasting had been required this would have increased the
phosphorus removal.
The sludge precipitated in the primary clarifier was a mixture of a cal-
cium phosphate compound, Ca(Ol-I) 2 , CaCO 3 , and coagulated organics plus grit.
This sludge was free flowing and considerably different from the gelatinous type sludge
produced in bench-scale studies. The sludge concentrated readily to 10 to 12% solids
by gravity thickening but was very difficult to dewater. The sludge removed from the
primary unit had a solids content of 6 to 8%. The investigators concluded that there
would be no practical value in attempting to digest the sludge derived from the lime
treatment process, inasmuch as biological activity of the sludge is suspended by the
high lime content. The sludge is free from noxious odors even after prolonged holding
under anaerobic conditions. The authors suggest that the sludge might be dewatered by
polymer conditioning and vacuum filtration or centrifugation followed by landfill
disposal.
The study results indicated that the organic capacity of an existing plant could be
increased without jeopardizing good treatment through the addition of lime to the
primary units. Laboratory studies using a combined flocculation-sedimentation system
with high rates of primary sludge return showed that recirculation of the lime sludge
increased the efficiency of the lime and improved effluent clarification. Supernatant
from the chemical sludge, even after 5 weeks of storage, contained no appreciable
amount of phosphorus, either total or soluble.
it was concluded that the lime treatment process could be readily controlled in
response to flow or on the basis of pH in a flash mixer or in the primary effluent. It
was also concluded that lime control in this manner would be simpler than that of a
chemical dependent on incoming phosphorus concentration.
5.3 Example of a Plant Design
During the period of January to April 1967, treatability studies for phosphorus
removal were made on the in fluent to the Rochester, New York treatment plant (4). At
this time the raw wastewater phosphorus concentration averaged about 8 mg/I as P.
Studies showed that the activated sludge process removed about 20% of the intluent
phosphorus. This removal agreed reasonably well with the phosphorus uptake normally
experienced with biological metabolism.
Jar tests were performed to determine the chemical treatment necessary to achieve the
State requirements of 80% phosphorus removal. Lime, and ferric chloride with anionic
polymers were used in the jar tests. Either 100 mg/I of lime as CaO (p 1 -I 9.5) or the
combination of 40 mg/I FeCl 3 and 0.5 mgf I Dow A23 was found to reduce the
phosphorus concentration of the supernatant by about 70%. Adding the removal from
the biological treatment would give an overall phosphorus removal of at least 80%.
BOD and suspended solids reductions indicated by these jar tests were 50% and 80 to
90%, respectively. Lime was selected as the chemical to be used.
5-3
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The present treatment plant facilities consist of settling tanks and lmhoff tanks, sludge
holding tanks, coil filters, arid sludge incinerators.
Expansion plans call for modification of present settling basins with flocculation
sections, addition of settling basins incorporating flocculation sections, conversion of
the lmhoff tanks to complete-mix activated sludge units, chlorination, and discharge to
Lake Ontario through a new lOft diameter 18,000 ft outfall line. Detention time in
the flocculation sections of existing settlers will be 20 minutes. In the new basins, the
flocculation time will be 40 minutes. Existing and new settlers will have overflow rates
of 1,000 gpd/1t 2 and detention times of 2.3 and 2 hours, respectively, at the average
design flow. A schematic diagram of the expanded plant is shown on Figure 5-1. Plant
design is based on the following critena
Average design flow, mgd 100
Maximum flow, mgd 200
Minimum flow, mgd 50
BOD loading, lb/day 250,000
BOD concentration, mg/I 300
Suspended solids concentration, mg/I 300
Volatile suspended solids concentration, mg/I 240
Lime will be added just after the waste passes through the ccmminutors. The pH will
be monitored ahead of the flocculation basins and lime addition will be automatically
regulated to maintain a pH of 9.5.
Anionic polymer will be fed at a point just ahead of the flocculation basins when
necessary to improve settling. Provisions for returning solids from the primary basins to
the flocculation basins have been included to enhance the flocculation process.
Although lime treatment together with biological treatment should remove 80% of the
incoming phosphorus as required by the State, chemical facilities for addition of alum
to the aeration tanks will also be provided. This system of phosphorus removal may be
used alone when the lime system is not in operation or may be used to supplement
lime clarification for greater phosphorus removal.
Since 80 to 90% suspended solids removal is expected in the primary basins, the
quantity of primary sludge based only on the increased suspended solids removal will
be about 1-1/2 to 2 times the normal sludge quantities. The sludge produced as a
result of chemical precipitation must be added to this quantity. It is anticipated that,
including the chemical sludge, about three times the normal amount of sludge will be
produced with chemical treatment. The primary sludge quantities at design flow with
lime treatment are estimated to be 175 tons/day. Without lime treatment the primary
sludge quantities will be about 65 tons/day. Since the Load on the activated sludge
portion of the process will be reduced by about 50%, the quantity of waste activated
sludge, which is very difficult to dewater, will also be reduced proportionally. At
design loads with lime treatment, the waste activated sludge quantities are estimated to
be about 30 tons/day as compared to about 75 tons/day if lime treatment were not
used. The solids handling facilities have been designed for a capacity of 270 tons/day.
5-4
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FIGURE 5-I SCHEMATIC OF PROPOSED
AT ROCHESTER, NEW YORK
PHOSPHORUS REMOVAL PROCESS
I POLYMER
I (WHEN NEEDED)
{J
ALUM
(WHEN NEEDED)
pH
L - _Y
(WHEN REQUIRED)
pH MONITOR
DI SCHARGE
TO LAKE
-------
The sludges will be combined and thickened in gravity thickeners to a minimum
concentration of 5%. The thickeners were designed for a solids loading of 13 lb/day/ft 2 .
After thickening, the sludge will be dewatered on vacuum filters and incinerated by
multiple hearth furnaces. Incinerator ash, which will contain some CaO and other
calcium and phosphorus compounds, will be sluiced and pumped to two, 0.5 acre
lagoons, each having a water depth of 11 ft and operating in series. The phosphorus is
not expected to leach out because of the high pH in the lagoons. The effluent from
the lagoons will be pumped back to the head end of the plant. It is expected that the
lagoons will have to be emptied every 9 to 12 months.
5.4 Capital and Operating Costs
The cost of adding rapid mix and flocculation facilities to existing plants has been
estimated for 1, 10, and 100 mgd plants. These estimates should serve as a general
guideline where phosphorus is to be removed by lime addition to the primary settler.
Estimated costs for rapid mix facilities include new concrete basins, having a detention
time of 1 minute at the design flow, and new mixers, with shafts and impellers
constructed of stainless steel. Flocculation facilities consist of the addition of the
necessary baffle walls and mixers to provide a flocculation chamber, having a detention
time of 30 minutes, within existing circular settling basins. Estimated costs of these
additions are given below and include an allowance for the Contractor’s installation,
overhead, and profit, plus an additional allowance of 20% of the construction cost for
engineering and contingencies.
Plant Size, mgd
Unit 1 12 122
( Estimated Costs, $ )
Rapid Mix 1,900 11,300 71,100
Flocculation 55,700 229,700 855,200
Total 57,600 241,000 926,300
The cost per thousand gallons for adding single-stage lime treatment to an existing
plant was determined by amortizing the capital cost at 6% over 25 years and adding
the cost of chemicals. Operational expenses such as labor and power are not included.
Capital investment for adding rapid mixing and flocculation equipment are given above.
Investment figures for the lime storage and feed systems are given in Chapter 10,
Section 10.3.3. Lime costs per 1,000 gal. are based on a chemical cost of $20/ton
with 90% available CaO at a dosage of 150 mg/I as CaO. The equipment and chemical
costs ( /l,000 gal.) are given below:
5-6
-------
Plant Size, mgd
Item 1 10 100
(Cost, /1,000 gal. )
Equipment Amortization
Rapid mix and flocculator 1.2 0.5 0.2
Chemical storage and feed system 0.4 0 2 0.1
Chemical Cost IA 14 14
Total Cost 3 0 2 1 1 .7
Costs listed in the table above do not include an allowance for additional equipment
for sludge handling and disposal and corresponding operating costs. Since a consider-
able quantity of additional sludge is produced, this expense is likely to be substantial.
5-7
-------
5.5 References — Chapter 5
1. Albertson, 0. E., and Sherwood, R. J., “Phosphate Extraction Process”, JWPCF,
41.8, p 1467 (1969).
2. Schmid, L. A., and McKinney, R. E., “Phosphate Removal by a Lime-Biological
Scheme”, JWPCF, 41:7, p 1259 (1969).
3. Black, S. A., and Lewandowski, W., Phosphorus Removal by Lime Addition to a
Conventional Activated Sludge Plant, Division of Research Publication No. 36,
Ontario Water Resources Commission, I 35 St. Clair Avenue West, Toronto 7,
Ontario, Canada (November, 1969).
4. 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 (Feb
9-10, 1971).
5-8
-------
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 simple. 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/I with an accom-
panying reduction of BOD and solids to 15 mg/I. For an additional l to 2 per 1,000
gal., phosphorus can be reduced to I mg/i or less and BOD and solids to 10 or 12 mg/I.
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 tnckling filter.
6-I
-------
Chemical addition to primary clarifiers wilt 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 final 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 underfiow 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
Both alum and sodium aluminate have been tested for phosphorus removal in trickling
filter systems. Where alkalinity is low, alum may depress the pH too far to give good
results. In these cases sodium aluminate would be preferred over alum.
The Fairborn, Ohio Wastewater Treatment Plant was used in a study (5) of direct dosing
to a trickling filter with sodium aluminate. However, the maximum phosphorus removal
obtained was only 66%. This emphasizes the need for addition at other points of the
process in order to obtain high reductions in phosphorus. Data on phosphorus removals
obtained, with both a dosed and control trickling filter, are shown in Table 6-I.
6-2
-------
Table 6-1
FAIRBORN, OHIO - TRICKLING FILTER PHOSPHORUS REMOVAL
Al:P weight ratio of I : 1 in dosed filter
Process Stream Total Phosphorus (mg/I) % P Removed
Range Average
Raw wastewater 9.2-18.0 13.2 —
Primary effluent 9.8-17.0 12.5 5.3
Final effluent, dosed 2.8- 6.2 4.5 66.0
Final effluent, control 9.0-12.8 10.9 17.4
Al:P weight ratio of 1.6:1 in dosed filter
Process Stream Total Phosphorus (mg/I) % P Removed
Range Average
Raw wastewater 10.2-36.0 14.8 —
Primary effluent 7.7-12.8 11.3 23.6
Final effluent, dosed 4.6- 8.4 5.2 64.8
Final effluent, control 8.0-14.2 10.6 28.4
Sludge analyses from the sodium aluminate studies revealed a volatile content of about
70% for the control filter sludge and 50% for the dosed filter sludge. The combined
sludge from one dosed and two normally operating filters was digested anaerobically. As
a result of only one of the three trickling filters being dosed, the excessive inert sludge
bulk reported by Dean (6) did not occur in the digesters.
The addition of alum to trickling filter effluent before final settling is being studied at
Richardson, Texas (7) (addition of iron and polymer at Richardson is covered later). A
plant schematic is shown on Figure 6-1 along with the points of chemical addition. At
design capacity of 1.6 mgd the treatment units have the following design
characteristics:
Pnmaiy Clarifiers
Detention time, hours 3.9
Overflow rate, gpd/ft 2 424
Tnckling Filters*
Hydraulic loading, gpd/ft 2 95
Process loading**, lb BOD 5 /1,000 ft 3 11.9
Depth, ft 6.5
Recirculation, gpd 70,000
Final Clan fier
Detention time, hours 2.6
Overflow rate, gpd/ft 2 416
* average
** assuming 30% BOD removal in primary
6-3
-------
FINAL
DIGESTED
_•j• SLUDGE
f -DRYING
BEDS
FIGURE 6—I
TREATMENT
RICHARDSON, TEXAS
PLANT FACILITIES
NAL CLARIFIER
FINAL SLUDGE
&
REC I
FILTERS
PRIMARY
CLARI FIERS
&
DI GESTERS
(UNHEATED)
RAW
WASTEWATER
TREATMENT
6-4
-------
Sludge is digested in the lower compartment of each primary clarifier prior to
dewatering on sand beds. Underdrain filtrate from the beds and digester supernatant,
after treatment with alum, are returned to the head of the plant. Major modifications
to the plant included chemical storage and handling facilities and a flash mix unit for
the junction box after the trickling filters. Chemicals are also added and mixed in the
raw wastewater wetwell.
Since the study will not be completed until early in 1972, the data and inferences
presented here are preliminary and not the final results. Influent phosphorus has been
about 8 mg/I (110 lb/day) of which about 70% is soluble. An Al:P weight ratio of
1 .83: 1 (molar ratio 2.1:1), fed exclusively after the trickling filters, reduced the
effluent phosphorus concentration to 0.5 mg/i
A split feed of 20% to the raw wastewater and 80% after the trickling filter at an
Al P ratio of 1.74:1 (molar ratio 2:1) reduced effluent phosphorus to 0.3 mg/I.
6.4 Performance Data Using Iron Salts
Addition of iron to trickling filters has been studied at Detroit and Wyoming, Michigan
and at Richardson, Texas
The Detroit Metro Water Department (8) 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 2 . The raw wastewater inorganic phosphate averaged 7.9 mg/I P with a range of
I .9 mg/I to 26.6 mg/I. Without the addition of chemicals, total phosphate removals
ranged from 18 to 57%. With the addition of 15 mg/I of iron (pickle liquor) removals
were 34 to 77%. With 15 mg/I iron and 0.3 mg/I Dow Purifloc A23 polymer, removals
ranged from 47 to 80%. With IS mg/I iron, 0.3 mg/I A23 polymer, and 30 mg/I
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 (9, 10, II). 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
FeCl 3 addition between the raw wastewater and the trickling filter effluent. The
chemical feed and phosphorus removal data are presented in Table 6-2.
6-5
-------
Table 6-2
IRON ADDITION FOR PHOSPHORUS REMOVAL AT WYOMING, MICHIGAN
Test No .
I .
Chemical Feed
Fe 3 ’ to raw wastewater, mg/i 15 15 15-20 IS
Fe 3 to trickling filter effluent, mg/I 5-7 3-4 0 7
Purifloc A23, mg/i, before primary 0.4 0.3 0.2-0.3 0
sedimentation
Total Phosphorus
Raw, mg/i 10.4 11.1 14.0 17.5
Primary effluent, mg/I 1.68 2 10 3.4 6.5
Removal, % 83.9 81.0 75.7 62.9
Final effluent, mg/i 0.88 1.46 2.9 5.7
Removal, % 91.7 87.6 79.2 67.4
To achieve the 80% phosphorus removal required by the State, iron addition of IS
mg/I to the raw wastewater and 3 to 4 mg/I Fe 3 to the trickling filter effluent
along with 0.3 mg/I 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 fiiter plant permitted greater than 80% phosphorus removal only with the
split chemical addition.
Initial data on iron addition at Richardson, Texas are included as examples in
following discussions on trickling filter plant dosage and iron leakage.
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,
fuli-scale trials may not be any more expensive than pilot-scale tests and they will be
more reliabie.
The jar test is a most important tool for both designer and operator when using
chemical addition Information can be obtained iiiexpensively 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
piant. 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,
6-6
-------
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 (12, 13) 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 (14),
liquid forms are more convenient to handle, especially at small plants. They may have
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 nietal 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.
6-7
-------
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 lb-mole/day by dividing by the atomic
weight (lb/lb-mole). Atomic weights are 31 for phosphorus, 27 for aluminum and 56
for iron. Next set the mole ratio (metalP) 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 lb/day,
310/31 = 10 lb-moles/day
If desired mole ratio (M:P) is 2:1,
( ) (10) = 20 lb-moles metal required
Using liquid alum having 4.37% Al,
(20) (27) = 540 lb Al required
(540)/(11.1 lb/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 lb/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 (15).
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 6-2. These curves can then be used to choose
the mole ratio required to meet effluent phosphorus standards.
6-8
-------
7
I I I
0
I I
I I I I
TREATMENT OF
PLANT INFLOW
0
0
0
0
0
0
I I I I
1 2 1 2
MOLE RATIO, Fe 3 /P
FIGURE 6-2 PHOSPHORUS REMOVAL BY IRON ADDITION (RICHARDSON, TEXAS)
6
5
TREATMENT IN
FINAL CLARIFIER
0\
(a,
I-
w
-J
L a-
LI
LU
CO
a
=
0
CO
a
=
a-
-J
I-
0
0
0
0
3—
2—
1
0
0
0
0
0
0
0
0
0
0
Oo
0
-------
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 I
mg/I or less. One direct way to judge polymer effectiveness is to observe effluent
turbidity. Polymer feeding can also be automated.
6.8 Iron Leakage
Iron leakage may be a problem in trickling filter plants using iron precipitation. Figure
6-3 shows some typical data from Richardson, Texas (7). 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 6-3 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/I as CaCO 3 . 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 floc 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 - 10
-------
(a) IRON LEAKAGE IN PLANT EFFLUENT
10 —
. S
E
—
WHEN TREATING IN
6 _ AL INFLOW
CLARIFIER
WHEN TREATING
S
-J
.
S t ’ 1 — r
0
0 I I I I I I
0.8 1.0 1.2 1. 1.6 1.8 2.0 2.2 2.1
MOLE RATIO, Fe 3 /P
6
(b) IRON LEAKAGE WHEN TREATING PLANT INFLOW
S
5—
IRON LOST IN
E
EFFLUENT
W/0 POLYMER
U i
-
-J
U.. S
LL
EF F LU EN T
2 IRON LOST
-J
o W/POLYMER IN FINAL
1 CLARIFIER
0 I I I I I I I
0.8 1.0 1.2 1.U 1.6 1.8 2.0 2.2 2.1
MOLE RATIO, Fe 3 /P
FIGURE 6-3 IRON LEAKAGE DURING PHOSPHORUS REMOVAL
611
-------
ANALYSES FOR CONVENTIONAL TREATMENT
R P F F R S S
A R I I E I U
W I I N C U P
H T A I D E
I R G R
E E
E C E N
F F A
F
F F
F T
FLOW
X
X
X
X
TOTSOL
X
X
X
X
X
X
X
TOTVOLSOL
X
X
X
X
X
X
X
SUSSOL
X
X
X
X
X
X
SUSVOLSOL
X
X
X
X
X
X
SET SOL
X
X
X
X
BOD
X
X
X
X
DO
X
X
COD
X
X
X
X
X
X
pH
X
X
X
X
X
X
TEMP
X
X
X
X
X
X
—
ADDITIONAL ANALYSES FOR CHEMICAL TREATMENT
R P F F R S S
A R I I E L U
W I L N C U P
H T A I D E
L R 6 R
E E C E N
F F E A
F F F T
F
PHOS
X
X
X
X
X
X
ALK
X
K
K
X
X
FE
K
K
X
X
X
X
AL
X
K
X
K
K
K
Soil
K
K
K
K
X
K
CL
X
K
X
TURB
K
X
X
TABLE 6-3 DESIRABLE LABORATORY ANALYSES
6 - 12
-------
6.10 References — Chapter 6
1. Nesbitt, J. B., “Phosphorus Removal — The State-of-the-Art”, JWPCF, 41:5, p
701 (1969).
2. HaIl, 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. Barth, E. F., Jackson, B. H., Lewis, R. F., and Brenner, R. C., “Phosphorus
Removal from Wastewater by Direct Dosing of Aluminate to a Trickling Filter”,
JWPCF, 41:11, p 1932 (1969).
6. Dean, R. B., “Reuse and Disposal of Lime and Alum Sludges”, presented at
Nutrient Removal and Advanced Waste Treatment Symposium, Cincinnati, Ohio
(Apr. 29-30, 1969).
7. Laughlin, J., “Modification of a Trickling Filter Plant to Allow Chemical
Precipitation”, presented at Advanced Waste Treatment and Water Reuse
Symposium, EPA, Dallas, Texas (Jan. 12-14, 1971).
8. Detroit Metro Water Department, “Development of Phosphate Removal Processes”,
Detroit, Michigan (July, 1970) (Program 17010 FAH Grant WPRD 5 1-01-67
Advanced Waste Treatment Laboratory, Cincinnati, Ohio), Pre-publication copy.
9. 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).
10. “Report on Wastewater Treatment Plant Improvements, Wyoming, Michigan”, Black
and Veatch of Michigan, Consulting Engineers, Kansas City, Missouri (1970).
11. Stonebrook, W. J., Dykhuizen, V., Beeghly, J. H., and Pawlak, T. J., “Phosphorus
Removal at a Trickling Filter Plant, Wyoming, Michigan”, Unpublished (undated).
12. TeKippe, R. J. and Ham, R. K., “Coagulation Testing: A Comparison of
Techniques — Part I”, JAWWA, 62:9, p 594 (1970).
13. TeKippe, R. J. and Ham, R. K., “Coagulation Testing: A Comparison of
Techniques — Part 2”, JAWWA, 62:10, p 620 (1970).
14. “Coagulants for Waste Water Treatment”, Chem. Eng. Frog, 66: 1, p 36 (1 970).
15. McAchran, G. E. and Hogue, R. D., “Phosphate Removal from Municipal
Sewage”, Wat. & Sew. Works, 118:2, p 36 (1971).
6 - 13
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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 phosthate 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 existLng facilities or estimated by knowledge of the
net solids production and phosphorus concentration in the waste activated sludge. Figure
7-I 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 1 938 (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), Barth
and Ettinger (5), and Eberhardt and Nesbitt (6). Barth, 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 (II) 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.
a
Lu
a
a
0
C ’,
a
-J
0
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a
Lu
a
Lu
C ,,
=
C d )
Ld
-J
I-
-J
a
I-
Lu
PHOSPHORUS REMOVED, mg/I as P
FIGURE 7 -I PHOSPHORUS REMOVAL CAPABILITY OF ACTIVATED
SLUDGE SYSTEMS RECEIVING SETTLED DOMESTIC WASTEWATER
0.
0.3
a
Lu
a
uJ
a
a
0. I
0
2
3
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
(IS). 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 nght.
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) arc
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/I as P for the raw wastewater and final effluent,
and less than 0.1 mg/I 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
-------
I Fe 3 1 7 • •j _
PH 413+17 0
6 9 6 8 6.8 6.6
6 7 6 6.0 5 t
IR AW WASTEWATE ]
FERRIC CHLORIDE
(43 Fe 3
FIGURE 7-2 TOTAL
SOLUBLE PHOSPHORUS
METAL ION DOSE
(Reprinted with permission from the American Institute
of Chemical Engineers, Mew York, M.Y. 10017)
RESIDUAL VERSUS
20
n
C
C O
=
C l )
=
uJ
C
C O
—j
0 10 20
30
METAL IO DOSE, mg/i
50 60
7-4
-------
1.2 — _______________
1.0 —
0 j 0 2 0 3 0 O 5.0
MOLES ALUMINUM ADDED
MOLE INITIAL TOTAL SOLUBLE PHOSPHORUS
FIGURE 7-3 INFLUENCE OF POINT OF ADDITION
ON PHOSPHORUS REMOVAL WITH ALUMINUM
(Reprinted with permission from the American Institute
of Chemical Engineers, New Ybrk, N.Y. 10017)
I I I
AL EFFLUENT
W WASTEWATER
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 I : I, 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 effcctivc. 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
74, 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-I 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 FeCI 3 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 carned
out at a pH of 6, only 20 to 40 mg/I of residual CaCO 3 buffering alkalinity may
remain (8). The bottom part of Table 7-I shows a companson 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
-------
I
I
20
I
AEROBIC BIOLOGICAL
TREATMENT ZONE
p
I
I ——-
PHOSPHORUS
REI4OVAL NITRIFICATIO
Ic
I
NO
METAL
ION ADDED
0
pH
FIGURE 7-U EFFECT OF pH ON PHOSPHORUS REMOVAL FROM A
FINAL EFFLUENT
(Reprinted with permission from the American Institute
of ChemicaT Engineers, New York, N.Y. 10017)
U)
c#)
a
0
U,
0
=
U i
-j
-J
a
(I )
I-
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I-
5
6
7 6 9
pH
a
L ii
0
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a
=
a-
UJ
-j
-I
a
-J
C -,
Ui
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a
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a
a
a
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I-
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L ii
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x
S
6
7
8
9
10
7-7
-------
Table 7-1
METAL SALT COMPARISON
Metal Salt
Liquid
Sodium Liquid Ferric
Aluminate Alum Chloride
(Al 3 ) (Al 3 ) (Fe 3 i ’)
(per lb of Metal Jon)
Bulk Delivered Price at Manassas, $ 0.69 0.32 0.23
Extraneous Ion Na+ SO 4 2 C1
introduced, lb 0.85 5.4 1.9
Maximum Alkalinity Loss as CaCO 3 , Lb None 5.5 2.7
(Per lb of Phosphorus at a Mole Dosage of 1: I)
Bulk Delivered Price at Manassas, $ 0.60 0.28 0.41
Extraneous Ion Na+ 504 C1
Introduced, lb 0.74 4.7 3.4
Maximum Alkalinity Loss as CaCO 3 , lb None 4.8 4.8
Phosphorus Removal of Metal Ion, lb 0.87 0.87 1 .8
Add itional Solids Production, lb AIPO 4 AIPO 4 FePO 4
3.9 3.9 4.9
Several other points should be considered in mineral selection. First, in the process
effluent the soluble Al, less than 0.5 mg/I (16), is apparently much lower than found
with iron additions, 6 mg/I Fe reported at Cleveland (19). Second, alum adds a
divalent ion (S0 4 2 ) 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/I range. An average Al:P
weight ratio of 2.1:1 was required to reduce the average influent total phosphorus of
3.6 mg/I to 0.45 mg/I 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 wilJ be much higher. Water treatment experience has shown
that velocity gradients greater than 75 ft/second/ft will result in the onset of chemical
floc disintegration (22). Undoubtedly, some chemical floc as well as biological floc
disintegration does occur in activated sludge tanks and, because of solids recycle, is
unavoidable. Floc 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 I I (8).
7.3 Performance and Optimization
7.3.1 Phosphorus Removal
At molar dosages of I to 2, metal ion to phosphorus, soluble phosphorus residuals of
0.3 to 3.0 mg/I 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) 1 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 scver.il (Yl)cs of
activated sludge systems were investigated, indicated that the most elfective total
phosphate reduction was provided by the conventional activated ‘dudge process with
iron (FeCI 2 derived from pickle liquor) fed to the primary effluent (23) The majority
of the test runs exhibited removals varying from 75 to 85% for inoiganic phosPhate
concentrations ranging from 1.9 to 26.6 mg/I as P in the raw wastewatci 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 ihie oiiveiitional
activated sludge process provided a treatment system capable of consistenily high
phosphorus reduction. The step-feed, activ4ted sludge piocess, with the .iddutiun ol iron
or iron, polymer and caustic soda, also provided consistent total phosphorus ictiuction
It was concluded that the full-scale plant should be designed for the activated sludge
process, arranged tor both conventional and step-feed configurations, and that provision
be made for phosphorus removal, utilizing injection of steel pickle liquor (FeCI . ,)
7-9
-------
A pilot plant (5,760 gpd) at Warren, Michigan (24) was designed to simulate the existing 36
rngd activated sludge plant, with the addition of one sand filter and one multimedia
filter operated in parallel alter 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 I .5
mg/I of Fe to 1.0 mg/I of soluble P0 4 3 reduced the average total phosphorus
concentration from 7 0 mg/I to 1.2 mg/I in the effluent. This amounted to a
reduction of 83%.
Leary, et at., of the Milwaukee Sewerage Commission (25), conducted a one-year
plant-scale study of the use of waste pickle liquor (FeSO 4 ) 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 1 15 rngd East plant durmg the test period. The 85 mgd
West plant served as a control. Pickle liquor feed was equivalent to an average Fe 2 ’
dosage of 9.4 mg/I. The pickle liquor was added before the aeration tank about
55 feet downstream from the point of addition ot return sludge. Based on a 1970
average total phosphorus conccntration of 8.2 mg/I 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/I at the East
plant and 1.4 mg/t 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%.
The pickle liquor addition increased the East plant effluent S0 4 2 concentration from
123 to 145 mg/I and decreased the alkalinity from 2 13 to 169 mg/I as CaCO 3 .
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) rdnged froni 2 I to 9.3%
H 2 S0 4 .
Comparison of the BOO, 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
Al: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 soIids, compared to a value of 80% in the control. The aluminate units had
higher p1-I, but all pH values were in the range of 7.7 to 8.2.
7 - 10
-------
5 I T I L
C-
U,
0l
e
a
C-
U)
a
C.
-j
I —
a
1
0 I I I I
0. 1$ 0.6 0.8 1.0 1.2 1.1$ 1.6 1.8 2.0
Al/P MOLE RATIO
FIGURE 7-5 COMPARISON OF THE EFFECTIVENESS OF ALUM AND SODIUM ALUMINATE
ON PHOSPHORUS REMOVAL
(Reprinted with permission from Journal Water Pollution Control Federation ,
in press, Washington, D.C. 20016)
a
C
SODIUM ALUMINATE
U
a
a
-------
Full-scale studies were conducted at Pennsylvania State University on the removal of
phosphorus by addition of aluminum to a 2.0 rngd activated sludge plant. These studies
began in 1969 and continued through 1970 (9). In the first phase of the work, alum
and sodium aluniinate 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 higj1er
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 tower 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/I 134 19 94 25
BOD 5 , mg/I 87 6 50 7
COD, mg/i 244 31 160 36
Phosphorus as P, mg/I
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 CaCO 3 , mg/I 156 103
pH 7.4 7.1 7J5 7.25
Alum Dosage as A1 2 (S0 4 ) 3 . 14H 2 0, mg/I 160
Na. Al 2 O 4 dosage, mg/I
A1.PRatio**,we ight 2.04:1 2.04:1
S0 4 2 , mg/I 24 131
*Chem lcals added at the effluent end of the aeration tank.
**Al.p ratio based on expression of phosphorus in the filtered ortho form.
7 - 12
-------
Aluminate addition was found to be most effective at a point in the aeration basin
near the outlet. Alummate 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 Additlon* Alum Addition
Suspended Solids, mg/I 110 22 26
BOD 5 ,mg/l 7! 9 13
COD, mg/I 172 55 68
Phosphorus asP, mg/I
Filtered Ortho 6.7 0.28 6 7
Filtered Total 6.3 0.36 6.7
Unfiltered Ortho 8.9 1.15 72
Unfiltered Total 10.0 1 41 7.3
Alkalinity as CaCO 3 , mg/I 168 80 120
pH 7.6 6.75 7 35
A1.P Ratio**, Weight 2.971
*Alum added to effluent of aeration tank. Values shown in this column are means
for the peziod of study.
**The Al:P ratio is based on expression of phosphorus in the filtered ortho iorm
Problems were experienced with loss of solids from the final basin during peak ilow
periods where alum treatment was being practiced. It was believed, however, iiiat
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 Dictiict 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/I as CaCO 3 ) at the fourth stage
of aeration using an Al:? weight ratio of 2:1 (molar ratio 2.3:1), based on an influent
phosphorus concentration of 8.15 mg/I as P. Mixed liquor suspended solids averaged
6,300 mg/I with a volatile content of 60% and a cell retention time of about 5 days.
Total phosphorus residuals averaged 1 .9 mg/I 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
p1-I 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
peiformance 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 PERFO RMANCE
Percent Less Than
Item Mm. 10 20 50 80 90 Max .
Primary Effluent
COD, mg/I 27 64 86 152 227 250 324
P,mg/I 2.2 38 50 8.4 1.4 12.8 18.5
BOD 5 :COD* 0.26 0.43 0.56
High Rate Activated Sludge Effluent
Soluble Phase
COD, mg/I 13 23 26 33 41 48 79
P, mg/i 0.2 0.8 1 .0 2.0 3.0 3.6 6.8
BOD 5 :COD* 0.14 0.28 0.36
Solid Phase
Suspended Solids, mg/I 0 10 14 21 36 44 196
BOD 5 .SS* 0.10 0.23 031
COD:SS* 0.43 0.58 0.80
P:SS* 0.039 0.05 1 0.063
Suspended Solids After
Polymer Addition,** mg/I 0 0 I 2 8 14 18
Suspended Solids After
Mixed Media Filtration**, mg/I 0 0 0 0 2 5 12
* Ratio of minimum, median, and maximum monthly averages for each parameter
**94 determinations
7 - 14
-------
1.2
DAILY POINTS, USING ALUMINATE
1.0 • 25 DAY MONTHLY AVERAGES,
USING ALUM
[ CHEMICALLY1 rCHEMICALLY1 rT0TAL SOLUBLE
REMOVED I = I REMOVABLE — I PHOSPHORUS IN
0.8 LPHOSPHORUSJ LPHOSPHORUSJ _SYSTEM EFFLUENT —
[ CHEMICALLY1 rPRIMARY 1 BACKGROUND
I REMOVABLE I I EFFLUENT I — OR
.•. [ PHOSPHORUS] I TOTAL I BIOLOGICAL
[ PHOSPHORUS] PHOSPHORUS
0.6 • REMOVAL —
LU
0
I pH = 7.2 to 7.8
— 0
0
0
>--J O
• 0
_I - 0 0 O
2 -
I I
0.5 1.0 1.5 2.0 2.5 3.0 3.5
ALUMINUM ADDED lb
CHEMICALLY REMOVABLE PHOSPHORUS ’ lb
FIGURE 7-6 CHEMICAL PHOSPHORUS REMOVAL ENVELOPE:
HIGH RATE ACTIVATED SLUDGE SYSTEM
-------
TO DIGESTER
;S SIAGE
¼ TI
FIGURE 7-7 0.2 mqd BIOLOGICAL NITRIFICATION - DENITRIFICATION PILOT PLANT
Greater Manassas Sanitary District
Prince William County, Virginia
EFFLUENT
FLASH MIXER
SEWAGE
SLUDGE
2ND STAGE
SEWAGE DIFFUSED
FROM AERATION
PLANT
WET WELL.
-------
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/I as CaCO 3 ) 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/I. Hais, et al. (27), indicated, using step aeration with alum addition,
that above p1-I 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/i effluent phosphorus residual) was achieved b ’ maintaining a final stage pH
between 6.3 and 6.6 at an Al P weight ratio of 2.1. Mulbargei (8) showed the
capability of pH adjustment with a denitrification system and demonstrated the
advantages of split chemical trcatmcnt in a multi-sludge system Results from the p1-I
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/I 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 p1-I, 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 cognizaril of
the phosphorus levels in the effluent suspended solids. There was a nominal 1 .0 mg/I
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/I of anionic polymer (Nalco
676), effluent suspended solids were reduced to 2 mg/I. Consequently, the effluent
total phosphorus was also reduced. Polymer performance was also reported (0 be
influenced by the pH (8). If total phosphorus residuals of less than 0.5 mg/I as I ’ are
required, a multimedia filtration system is recommended. Dual and (ri-media filtration
both provide effective tertiary solids separation. Hais. et al (27) showed Lh ,.iF tn-media
filtration consistently removed between 5 and 10% more uspciidcd u,’ imcrial than did
the dual-media filter. Filter runs between 24 and 2 hi..uiis vt:r maintamimed
Tn—media filtration performance was repc’wted to i . si.’ hicriot to . ‘‘-‘‘‘‘‘a iRi ,itton
at Manassas (8).
7 - 17
-------
.0
C ,)
_ -
o -
=
Q
0
=
Q_ ..—
a’
LU. E 0.5
-J
0
_J C,)
LU
I-
0
0
ALUMINUM ADDED lb
CHEMICALLY REMOVABLE PHOSPHORUS ’ lb
FIGURE 7-8 EFFECT OF ALUMINUM ADDITION WITH pH
ADJUSTMENT ON TOTAL SOLUBLE PHOSPHORUS RESIDUAL
a’
E
• a
C,) —
a
— 00
E
LU
0
—
—
- _J
LU
a
a a
LIJLU
LU
C,,
a
Wa
-J
C )
C-)
a
a
ALUMINUM ADDED, mg/i as AI 3
FIGURE 7-9 -CHANGE IN SLUDGE VOLUME INDEX
AND ADDITIONAL SOLIDS PRODUCTION DUE
TO ALUMINUM ADDITION
(Reprinted with permission from the American Institute
of Chemical Engineers, New York, N.Y. 10017)
3.0 ‘LO 5.0
1.0 2.0
0
00
0 5 tO 15 20 25
7 - 18
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7.3.2 Sludge Characteristics
Long, et al. (9), indicated in studies conducted at Pennsylvania State University that
sludge production in the chemical-biological process was significantly greater than in
biological treatment only. This is to be expected because of the sludge produced by
mineral addition. For a 1 year alum test period, the amount of waste sludge from the
chemical-biological process was 2.1 times by weight that from the biological process
alone. On the basis of volatile solids, the ratio was 1.56:1. It should be pointed out,
however, that high alum doses (mole ratios of Al:P of 2:1 or above) were used in the
study and that the sludge age of the undosed control was greater than that for the
dosed system
Figure 7-9 shows the changes found at Manassas (16) in sludge density and additional
solids production with aluminate additions. During these studies the average weekly
primary effluent phosphorus concentration ranged between 9.3 and 13.1 mg/I. From
this figure it can be concluded that up to stoichiometric requirements, additional solids
production will be about 4 lb/lb A1 3 +. Above stoichiometric requirements, production
of additional solids will drop off. This observation can be explained by the
predominant formation of AIPO 4 (molecular weight, 122) at less than stoichiometric
Al additions and formation of Al(OH) 3 (molecular weight, 78) above stoichiometric Al
additions. Figure 7-9 also shows an improvement in sludge density (which is inversely
related to the sludge volume index) with increasing aluminum dosages. Whether or not
this density improvement is enough to compensate for the additional solids that must
be handled, and thereby eliminate the need for additional sludge disposal facilities, is
dependent upon the original sludge density and the amount of phosphorus removed
and aluminum added. Figure 7 -1 0 shows actual plant data which indicate mineral
additions do produce a dense, stable sludge.
Figure 7-10 also shows the change of the sludge settling characteristics as a function of
the operating solids level (8). The reported sludge settling characteristics were
determined in I -liter graduates. Manassas experience also indicates that the settled
sludge does have a tendency to classify (separation of biological and chemical
fractions) although not to such an extent as to harm system performance. Therefore,
sludge collection systems which have a tendency to classify sludges into heavy and
light fractions, i.e., those that work with a hydrostatic or vacuum principle, probably
should be avoided.
Sludges resulting from aluminum addition are amenable to anaerobic digestion (28). In
conjunction with laboratory scale aluminum addition to activated sludge units at the
University of Missouri at Rolla (26), laboratory scale anaerobic digesters were studied
to determine the fate of the aluminum and phosphorus. The digesters were fed 70%
activated sludge and 30% primary sludge. Detention time in the digesters was 1 5 days.
The fate of the phosphorus in the digesters is shown in Figure 7-1 1. The phosphorus
which had been precipitated with aluminum was concentrated in the digester sludge
and not released to the supernatant. Tests confirmed that the aluminum was also
retained in the digester sludge. Comparing the cligesters which were fed aluminum
7 - 19
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MIXED LIQUOR SUSPENDED SOLIDS, mg/i
FIGURE 7-10 PHYSICAL CHARACTERISTICS OF
HIGH RATE ACTIVATED SLUDGE WITH MINERAL ADDITION
E
U i
Ui
x
-J
Ui
-J
U)
c’4
‘I-
-u
a.
Ui
I-
-J
I-
I-
Ui
U)
1000 2000 3000 ‘ O00 5000 6000
7 - 20
-------
I I I I I I I
1600 LEGEND
DIGESTER FEED SLUDGE Sodium Aluininafe Sludge
Alum Sludge
Control
-*
—— —
-
800 .—- - •— —— —
/ .—
/ -...---—
—‘ /
__.__‘ \ ,1/
4 ___ _‘ /
U
DIGESTER SLUDGE
u 1600
-
E - /
, —-—--
C ,, /__—-_-__ _-__
800-
C ,,
_I -
I-
0
I-
0
80 DIGESTER SUPERNATANT
-
‘k — -
-------
sludge with a control digester, using the parameters of volatile acids and gas
production, the digester performance was equivalent.
Experience at Manassas has found no resolubilization of phosphorus with the waste
organic-chemical sludge in anaerobic digestion (16). Connell reports similar observations
with iron additions (10).
At Manassas, comparative bench studies of settled high rate activated sludge
dewaterability with and without aluminum additions showed that dry cake filter yields
will substantially increase with aluminum additions, but there is a tendency to produce
a wetter cake. Similar high yield rate observations after aluminum additions have been
observed by Smith, et a!. (29). This increase in dewaterability can probably be
explained by the fact that with efficient chemical utilization, aluminum phosphate, not
aluminum hydroxide, is the inorganic precipitate being dewatered.
Mulbarger (8) indicates that it would be reasonably conservative to design solids
handling equipment for the organic solids loading on the basis of conventionally used
cnteria, assuming that the additional chemical solids that must be handled will be
compensated for by the increase in sludge dewaterability.
No published data are available on the dewatering characteristics of iron-organic
sludges.
Aluminum additions will result in the least amount of additional solids for ultimate
disposal, probably 30 to 50% of that found with iron and 5 to 20% of that found
with two-stage lime treatment.
Aluminum-organic sludges may be incinerated or spread on land. When the sludge is
applied to the soil, leachates should not be a problem because of the natural
phosphorus exchange capacity of soils and the low solubility of the aluminum
phosphate precipitate. Preliminary tests by Dotson (30) showed no plant-inhibitory
effect with applications of organic-aluminum sludges of up to 1 in. per week although
some initial seed germination inhibition may be encountered at higher applications.
Generally, it is felt that the soil building and fertilizing benefits derived from organic
sludge application more than compensate for any deleterious effects from the inorganic
aluminum sludge.
7.3.3 Costs
Power costs for combined chemical-biological treatment are negiigible. 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 phosphonis 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 - 22
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7.4 Process Design Examples
Four design example problems were selected br dli .11 bitrary wastewater Each
illustiates specihe points in combined chemical-biological wastewatci treatment with
alum additions for phosphorus removal Results 1 mm the Manassas work are used for
some of the calculations
Assumptions:
a Net biological solids production 0 28 lb of volatile solids produced per lb
of COD removed, the volatile solids contain :y; Pliolillol us, and the total
solids average 80% volatile
b The biological system is to be designed with .i miiiiniuni hydraulic contact
time of I hour at the maximum 4 how hydraulic peak. The ratio of
maximum flow tale to average flow rate = I 5 A cell retention time (CR1’)
of two days is selected
c Primary Efiluent Characteristics
COD = 200 mg/I
P = 10 mg/I
Alka ltnity = ISO mg/I as Ca(’0 1
Find:
I Ef lluent COD, BOD 5 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 tollowing delivered prices
Alum = $0 32/lb Al 3
Polymer = $ I 50/lb polymei
Acid = $0.02/lb H 1 S0 4
Lime = $001/lb Ca(OFl)-,
Example No. 1: High rate system for most efficient chemical addition with pH
above 7 and no optimization by pH adjustment
1 Effluent COD, SOD 5 , 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 )
7 - 23
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Effluent soluble COD = (4 1-33) + 33
= 5+3338mg/I
Assume effluent suspended solids 20 mg/I
Total effluent COD 20 (0.58) + 38
12 + 38
50 mg/I
Effluent BOD 5 = 38 (0.28) + 20 (0.23)
= 10.6 + 4.6 [ Using BOD 5 ratios
from Table 7-4]
= 15 mg/i
Background phosphorus removal per lb COD removed
= (biological solids production ratio) (phosphorus content)
= (0.28) (0.02)
= 0.0056 lb phosphorus removed
lb COD removed
[ The background phosphorus removal rate at Manassas was found to be approximately
0.01 lb P Removed 8 1
lb COD removed
Background phosphorus removal expressed as concentration
= (COD removal) (phosphorus removal per lb COD removed)
= (200-38) (0.0056)
= 0.9 mg/I
Chemically removable phosphorus = 10.0 - 0.9 9.1 mg/i
Use Al 3 l7chemIcaIIy removable phosphorus dose = I .25 lb/lb
From Figure 7-6 find 0.6 lb of phosphorus removed per lb of aluminum added at the
dosing ratio of I 25
Aluminum dose = I 25 (9.1) = 11.3 mg/I AJ 3
Chemical phosphorus removal = 0.6 (ii .3) = 6.8 mg/i P
Soluble phosphorus residual 9.1 - 6.8 = 2.3 mg/i
Total effluent phosphorus = 2.3 + 20 (0.05) EP/SS ratio from
= 3.3 mg/I P Table 7-41
7 - 24
-------
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. Total Solids Production
In calculating inorganic solids productions, dissolved solids additions, and
alkalinity losses derived from alum additions, the designer should be aware
of the two pH dependent competing reactions
a. A1 2 (S0 4 ) 3 + 6HC0 3 2Al(OU) + 3 SO + 6C0 2
(54 as AI 3 ) (300 as CaCO 3 ) (156) (288)
or
5.5 lb of CaCO 3 alkalinity lost per lb of Al 3
2.9 lb of Al(OH) 3 produced per lb of Al 3
5.4 lb of S0 4 2 added per lb of AI 3
b. Al 2 (S0 4 ) 3 + 2P0 4 3 2 AIPO 4 + 3S0 4 2
(54 as Al 3 ) (62 as P) (244)
or
4.5 lb of AIPO 4 produced per lb of Al 3
3 9 lb of AIPO 4 produced per lb of P
0.87 lb of Al 3 used per lb of P
Because of phosphorus removal and aluminum uptake by “sorption” mechanisms, these
preceding reactions do not completely describe observed phenomena but the estimates
derived from them are close enough for engineering design.
Solids production due to AIPO 4 formation
3 9 mg AIPO 4
(6.8 mg/I phosphorus removed chemically) ( )
mg P
= 27 mg/I AIPO 4
Aluminum used in AIPO 4 formation
= (6 8 mg/I P) ( V .0 7 mg Ai
mgP
= 5.9 mg/I AI 3
Aluminum available for hydrolysis
11.3 - 59 5.4 mg/I Al 3
7 - 25
-------
Solids production as Al(OH) 3
= 5.4 mg/I Al 3 (2.9 mg Al(OH) 3 ) = 16 mg/I
mg Al 3
Total inorganic solids production per lb Al 3 added
= 27 + 16 = 3 8 lb of solids
11.3 1bofAl 3 added
This result compares closely with the data in Figure 7-9. These calculations illustrate the
point brought out in the text; i.e., the magnitude of solids production is a function of
the proximity of aluminum dosage to the stoichiometric phosphorus requirement,
assuming the pH is in the proper range.
Biological solids production:
= (COD Removal) (Biological Solids Production Ratiok
Volatile Fraction
= (200-38) (0.28)
= 57 mg/I SS
Total Solids Production = Inorganic + Biological
=27 + 16+ 57
= 100 mg/I SS
3. Alkalinity Loss
Maximum possible alkalinity loss due to aluminum hydrolysis
= 5.4 (5.5) = 30 mg/i as CaCO 3
which gives = 2.7 lb CaCO 3 loss/lb AI 3
At an aluminum added to chemically removable phosphorus ratio of 1.16, an average
alkalinity loss of 2.3 lb CaCO 3 /lb A1 3 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.
4. Dissolved Solids Addition 2-
mgSO 4
SO ( = (11.3 mg Al /l) (5.4 ) = 61 mg/i
mgAl +
7 - 26
-------
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).
5. Final Clarifier and Pumping System
Organic solids = 57 mg/I
inorganic solids = 43 mg/I
Total = 100 mg/I
Solids lost in effluent = 20 mg/I
Solids accumulation = 80 mg/i
CRT = 2 days, and aeration tank detention at average flow
= 1.5 (1) = 1.5 hours
MLSS in aerator = ( 2 (2 ) (80) = 2,560 mg/I
Since this is the final clarifier a conservative overflow rate should be utilized. From
Figure 7-10 the minimum settling rate of 1,150 gpd/ ft 2 was observed at a MLSS of 2,600
mg/I. To provide for peak flow, the clarifier should be designed for an overflow rate of
770 gpdfft 2 , 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/I MLSS = 82 mI/gm (Figure 7-10)
Approximate return sludge concentration at worst condition
io6 106
= = 12,200 mg/I
Solids balance equation where R is sludge return rate
max + R) (MLSS)< (R) (return sludge conc.)
or considering just the case of equality,
R=l5Q 2,600
ave 12,200-2,600
R = 0.4 gave
(Actually, effective design would indicate a ±50% capability; i.e.,
0.2 to 0.6 Qave capability)
Solids to be wasted per day = 80 mg/l at a minimum sludge solids concentration = 12,200
mg/l. Assume continuous wastage with separate pump.
7 - 27
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= gave 12,200 = 0.0065 cave
(Again, effective design would indicate a ±50% capability; i.e., 0.003
to 0.01 Qave capability).
6. Alum Cost
= (11.3) (0.32) (8.34)(lo ) = 3.l /l,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. I but with sulfuric acid addition
for pH adjustment (residual alkalinity = 30 mg/I as CaCO 3 ) 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, BOD 5 , and suspended phosphorus concentrations in the effluent are
substantially reduced by an assumed polymer dose of 0.3 mg/l. The
chemically removed phosphorus was calculated for this case by multiplying the
0.6 lb P removed/lb Al 3 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/I with an aluminum dose of 11.3 mg/I. But
Figure 7-8 showed from experimental data a nominal 0.5 mg/I residual
phosphorus at this Al:P ratio and, therefore, the conservative residual of 0.5
mg/I soluble phosphorus was chosen.
2. Inorganic solids production consisted of 34 mg/I as AIPO 4 and 11 mg/I as
Al(OH) 3 due to the preferential formation of AIPO 4 at an optimized pH.
3. The alkalinity loss due to hydrolysis was also reduced due to increased AIPO 4
formation. Using 96% strength sulfuric acid and assuming a residual alkalinity
of 30 mg/I as CaCO 3 , a calculated sulfuric acid requirement of 101 mg/I was
found. The sequence of chemical addition is acid, alum, and polymer, with
about 5 minutes of mixing at peak flows between each application.
7 - 28
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Table 7-5
SUMMARY OF DESIGN CALCULATIONS
Item Example I 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/I 11.3 11.3 27.3 18.2
Effluent Quality, mg/I
Suspended Solids 20 2 0 0
COD 50 39 38
BOD 5 IS II 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/I
Biological 57 57 57 57+
Chemical 43 45 92 66
Alkalinity Loss
Due to Al 3 Addition 30 21 109 59
mg/I as CaCO 3
so 2- Increase Due
to Al + Addition, mg/I 61 61 147 98
Alum Cost, / 1,000 gal. 3.1 3.1 7.4 4.9
Polymer Cost, / 1,000 gal. 0 0.3 0.3 0.3
Acid Cost, / 1,000 gal. 0 1.7 0 1.0
Total Cost, / 1,000 gal. 3.1 5.1 7.7 6.2
Clarifier Overflow Rate, gpd/ft 2 770 670 360
Return Sludge Rate 0.4Q 0.6Q 0.9Q
Waste SludgeRate 0.0065Q 0.008Q 0.Ol2Q
*Spllt treatment here incorporates mineral addition as in Example No. I without pH
adjustment, followed by mmeral addition with pH adjustment at a later point.
4. Dissolved solids additions totaled 156 mg/I S0 4 2 ; 61 mg/l due to the alum and
95 mg/I due to the acid.
5. Because of improved solids capture, the operating MLSS in the aerator
increased to 3,200 mg/I necessitating a lower surface overflow rate in the
clarifier, and greater return and waste sludge handling capability.
7 - 29
-------
6. Chemical costs increased to 5.I /I,0O0 gal. with benefits in phosphorus,
suspended solids, BOD 5 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/i 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 Al:P weight ratio of 3 (Figure 7-8)
was used.
2. Inorganic solids .production consisted of 35 mg/l as AIPO 4 and 57 mg/I as
Al(OH) 3 . These values are approximately double the total inorganic solids
productions found in Examples No. I and 2.
3. The addition of alum caused an alkalinity consumption of 109 mg/i, which
would deplete the background alkalinity of the wastewater to 41 mg/I as CaCO 3 .
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.
4. Dissolved solids addition consisted of 147 mg/I SO 4 2 from the alum.
5. Additional solids productions resulted in an operating solids level of 4,800 mg/I
in the aerator, necessitating a lower surface overflow rate and higher return
and waste sludge capabilities than in earlier examples.
6. 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/I. Assume the
first point of addition is as in Example No. I. Further mineral addition will be
made to the effluent in Example No. I either before filtration or during a second
stage of biological treatment.
I. 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. I 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 3 + dose of 6.9 mg/I for the second addition was arrived at using Figure 7-8
for a pH adjusted system.
7 - 30
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2. Additional inorganic solids production was 8 mg/I as AIPO 4 and IS mg/i as
Al(OH) 3 .
3. The alkalinity depletion due to aluminum addition was calculated as 59 mg/I
as CaCO 3 . This would leave a residual alkalinity of 91 mg/I. Sixty mg/I as
H 2 S0 4 was added to reduce the alkalinity to 30 mg/I. Again, in a real
situation acid might not have to be added to this wastewater.
4. Dissolved solids additions were calculated as 98 mg/i SO 4 2 due to alum
addition and 59 mg/I SO 4 2 due to acid addition.
5. Clarifiers were not sized in this example. If two biological processes were used
in series, the final clarifier of the flrst process could be designed with an
overflow rate from the middle of the performance curve in Figure 7-10.
6. The cost of acid is l.0 / 1,000 gal. Total chemical costs are 6.2 / 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 - 31
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7.5 References — Chapter 5
I. Guggenheim Process Described by: Kiker, J. E., “Waste Disposal for Dairy Plants”,
Sanitarian (Los Angeles) 16, p ii to 17 (1953). Florida Engng. Exp. Station 7,
Leaflet Series No. 51 (1953). Moore, R. B., “Biochemical Treatment for Anderson,
md.”, Wat Works and Sew., 85 11(1938).
2. Thomas, E. A., “Phosphat-Elimiriation in der BelebtechlemanrilagevonMannedorf und
Phosphat-Fixation in See-und Klarschlamrn”, Viertelja/,rsschrift der Naturforschenden
Gesellschaft in Zurich, 110, Schlussheft, S., p 419 (1965).
3. Wirts, J. J., “Communications”, Buckeye Bulletin (Summer, 1966).
4. lenny, 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. 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, I 97 I).
9. Long, D. A., Nesbitt, J. B., and Kountz, R. R., “Soluble Phosphate Removal in the
Activated Sludge Process — A Iwo 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 (111) and Iron (III) Salts”, Water Pollution Control Research Series, 17010
EKI, Contract 14-12-158, USD1, FWQA (April, 1970).
7 - 32
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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)
IS. Cuip, R. L., and Cuip, 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 Frog. Symp Series. 67. 1 07, p 107
(1970).
17. Sawyer, C. N., Chemistry for Sanitary Engineers, MCraw-HiIl, New York (1960).
18. McKee, J. E., and Wolf, i-I. W., Waler 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, p32 (1971).
20. Water Quality and Treatment, 2nd Ed., American Water Works Association, New
York (1951).
21. Ockershausen, R. W., Harriger, R. D., and Zuern, I-I. E., “Phosphorus Removal Tests
with Alum” for Buffalo Sewer Authonty, Buffalo, New York, by Technical Service
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), Prepublication Copy (July, 1970).
24. Boggia, C., and E-Ierriman, 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. I IO1OFLQ, 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).
7 - 33
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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 Labonitory,
Office of Research and Monitoring, Environmental Protection Agency, Cincinnati,
Ohio (July, 1971).
7 - 34
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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(HCO 3 ) 2 + Ca(OH) 2 2 CaCO 3 + 2 H 2 0
NaHCO 3 + Ca(OH) 2 CaCO 3 + NaOH + H 2 0
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 3 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 3 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
Mg 2 + 2 OW ÷ Mg(OH) ,
This reaction does not approach completion until the pH is raised to II. Magnesium
hydroxide is a gelatinous precipitate which aids colloid removal, but which hinders
sludge thickening and dewatering. Where the pH must be raised to I I or above to
meet a phosphorus removal requirement or some other treatment objective, Mg(Ol-f) 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 p1-I and precipitate the
excess lime as CaCO 3 in the second stage. The following reaction occurs
Ca 2 + CO 2 + 2 0H ÷ CaCO 3 + H 2 0
Carbon dioxide may also be used to lower pH after lime treatment. The important
reaction in this case would be the conversion of Co 3 2 to HC0 3 .
8-I
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A final reaction that is of major concern, where lime is to be recovered by sludge
recalcination, is as follows:
CaCO 3 - - CaO + CO 2
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 II, 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 appropnate 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 CO 2 may be necessary before discharge and will
almost certainly be required before filtration to prevent post-precipitation of CaCO 3
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 merts 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, CaCO 3 , and
Mg(OH) occurs. Consideration of the solubility product for CaCO and the
equilibrium between CO 3 - and HC0 3 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
concentration of Ca 2 + present in the water. In two-stage treatment CO 2 is added after
the first-stage settler to bring the pH down to about 10 where CaCO 3 precipitation
results. The CaCO 3 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
CuIp and CuIp (I). Although the data on handling of wastewater sludges for
recalcinatiori and the recalcination process itself are limited mainly to this plant,
8-2
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WAS it
WASI1WATER
FILTER
RECYCLED LIME
00
MAKEUP
L I ME
WATER
CARBON
DIOX IDE
CAR BO N
DIOXIDE
C EN TRA I E
SLUDGE
WASKWATER
WASTE SLUDGE
LIME DISPOSAL
FIGURE 8-I SINGLE STAGE LIME TREATMENT SYSTEM
-------
WASTEWATER FEED
00
WASTE
WASh WATER
FILTER
WATER
SLUDGE
CARBON
DIOXIDE
LIME OR DISPOSAL
WASNWATER
SLUDGE TO RECALCINATOR
FIGURE 8-2 TWO STAGE LIME TREATMENT SYSTEM
-------
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 II. To reach that pH requires about 300 mg/I 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 (I to 20 mg/I) 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/I P. Before the filters the phosphorus concentration is
about 0.4 mg/I.
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. Cuip and CuIp report
typical filter effluents with a BUD of 3 mg/I, COD of 25 mg/l, 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 ca lcined 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 CaCO 3 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.
8-5
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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/i as CaO, and the pH after recarbonation was
maintained at about 10.3. Ferric iron was added at a rate of S mg/i 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 remov il for the system was similar to that obtained at Tahoe; 0.09 mg/I
as P was the average concentration remaining. Phosphorus concentration before
filtration averaged 0.13 mg/I.
In addition there was significant organic and suspended solids removal. The average
BOD was reduced from 15 mg/I to 2.1 mg/I before and 1.5 mg/I after filtration.
Suspended solids were reduced from 33 mg/I to 17 mg/I before filtration and 3.8
mg/I 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 I .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
recalcmed 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 I I 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 p1-I is clearly indicated. Since the average phosphorus concentration of the
secondary effluent was about 10 mg/I 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/I 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/I. This
was reduced to 16.5 mg/I 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
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2.0
CLARIFIER pH
FIGURE 8—3 EFFECT OF pH
ON PHOSPHORUS CONCENTRATION
OF EFFLUENT FROM FILTERS
FOLLOWING LIME CLARIFIER
E
I-
I—
z
0
C-)
C ,)
0
=
C ,)
0
=
1.5
1.0
0.5
0
8.5 9.0 9.5 10.0
10.5
11.0
11.5
8-7
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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 fdters. 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 CaCO 3 , 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-I. It appears from the data already available that the lime dose will usually be
in the range of 300 to 400 mg/i as CaO for two-stage treatment, and from 150 to
200 mg/i where single-stage treatment is satisfactory.
Table 8-I
LIME REQUIREMENTS
Feed Water Approximate Lime
Alkalinity Clarifier pH Dose
(mg/i as CaCO 3 ) (mg/I of CaO)
300 9.5 185
300 10.5 270
400 9.5 230
400 10.5 380
8-8
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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/I as CaCO 3 or less, form a poorly settleable floc because of the low fraction of
dense CaCO 3 . A pH above 10 is needed even to obtain measurable CaCO 3 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 CaCO 3 sludge in situations
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 solubiity 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 CaCO 3 sludge, prevents solids carryover when the first stage settler does not
operate properly.
With high alkalinity waters, a well settling floc 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
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removal, the second stage would be used in these cases to lower calcium content in
the effluent and to obtain high CaCO 3 sludge for lime recovery. There is not yet
available operating expenence at enough plants to state positively the alkalinity at
which single-stage treatment will perform satisfactorily, from the standpomt of floe
settleability. At present, experience indicates that at an alkalinity of I 50 mg/I as
CaCO 3 , single-stage treatment probably cannot be used. Between 150 mg/I and 200
mg/I settleability will probably depend on the amount of organic floe present Above
200 mg/I, 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. Culp and CuIp (I) 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 aiid
some design parameters are given by CuIp and Culp. 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 how of about one third of
design flow.) Mixing is accomplished with a vertical shaft mixer. The flocculator is a
8 - 10
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WASTEWATER LIME
[ 1 EFFLUENT
___ __ L
DO
SLUDGE
FIGURE 8-4 TYPICAL UPFLOW CLARIFIER
8 - 11
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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 2 . 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 84 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
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Tahoe. Whether modifications m 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
was operated at a constant overflow rate of I ,440 gpd/ft 2 . This proved satisfactory
even with a sludge blanket. At Washington, D. C., tests were made with settler rates as
high as 1,950 gpd/ft 2 although the average rate was 1,120 gpd/ft 2 . For design, a peak,
dry weather overflow rate of 1,200 to 1,400 gpd/ft 2 , 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 ne’cessary to make the effluent suitable for filtering or for
discharge. An excellent discussion of all aspects of wastewater recarbonation is given
by Cuip and Culp (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, CO 2 generators,
and commercial liquid CO 2 .
The equipment used for contacting the CO 2 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 CO 2 source is 8 ft. Pilot studies at’ Washington,
D. C., with a tank depth of 11 ft and a turbine mixer to reduce bubble size and
distribute bubbles, showed almost 100% absorption of CO 2 .
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 CO 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. Cuip and Cuip
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.
8 - 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 CaCO 3 . Where lime
recovery is practiced, maximum formation of CaCO 3 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 CaCO 3 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 CO 2 feeding equipment.
CuIp and CuIp discuss the calculations in detail. It was found in pilot plant work at
Washington, D. C. that the CO 2 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 CaCO
formation with the CO 3 - coming from the CO 2 . 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 CO 2
dose in mg/I is then equal to:
(Ca reduction in mg/I) ( )
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
minute detention time and a 2,400 gpd/ft 2 overflow rate at design flow. No
flocculation equipment is provided. CuIp and CuIp (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 uptiow equipment of the integrated type discussed
earlier. Without a flocculating aid, fine CaCO 3 precipitate escaped the settler, even
at a settler overflow rate as low as 1,100 gpd/ft 2 . Using 5 mg/I of Fe 3 , good
operation was observed at overflow rates as high as 1,950 gpd/ft 2 . Hydraulic
limitations prevented testing at higher rates. Apparently, the good contact between
solids and water during flocculation was beneficial in reducing CaCO 3 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 2 . 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 2 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 penods
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 tn-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 2 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 2 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 2 . 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 2 . Higher rates could result in inconveniently short runs during periods of high
solids load. Rates much lower than 3 gpm/ft 2 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 2 rate is
based upon having good operation of the second-stage settler. Frequent settler upsets
must be avoided.
Tn-media filters are used at the Tahoe plant. These are pressure filters which at design
flow operate at 5 gpm/ft 2 . 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 2 . The backwash water is reprocessed through the lime
treatment system. More details are given by CuIp and CuIp (1).
From the available data, it is difficult to give precise criteria for choosing between
dual-media or tn-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
floc weakness, the tn-media filter offers, therefore, more protection from turbidity
breakthrough. The cost, however, is slightly higher. Where the tn-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, tn-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 1 0 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 I .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/l,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 CaCO 3 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/I as Ca.
Design information for the Tahoe sludge handling system is reported by CuIp and
CuIp. (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 lb/day/ft 2 and the design overflow rate is 1 ,000 gpd/ft 2 . 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 underfiow 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 3 . 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
8- 17
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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 CuIp and CuIp. 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 I ,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
CaCO 3 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. TIns 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 l)laflt a
flexible hose was used to feed lime slurry. By periodically flattening or flexing the
hose, the scale was removed. The mechanical niixer 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.
S - 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 Cuip and Cuip (I). Capital cost for the 7.5 mgd lime treating facility was
$1,115,000 and cost of the filters was an additional $705,000. For 1969 the estimated
operating costs exclusive of equipment amortization were 7.3 /l ,000 gal. for lime
treatment without filtration and 2 . 8 ç /l,OOO gal for filtration. Amortization of
equipment, based on costs adjusted to the 1969 national average, interest of 5% for 25
years, and the assumption of the plant operating at full capacity, would add
2.7ctf 1,000 gal. for treatment without filtration and 1.8/l,000 gal. for filtration. For
the plant operating at full capacity the total cost of operation would then be
l0.0 /1 ,000 gal. without filtration and l4.6 /l ,000 gal. including filtration. The
fraction of the cost resulting from amortization is a substantial 31% even with the
assumption of full capacity. There is an obvious need to keep equipment size to a
minimum. Flow equalization deserves strong consideration when planning for lime
treatment of secondary effluent to minimize the initial cost of new equipment.
Additional costs for lime treatment based on information from Tahoe and other
sources have been reported by Smith and McMichael (6). Tables 8-2 and 8-3 show
Table 8-2
CAPITAL COST OF LIME TREATING FACILITIES
Cost ($)
Treatment Plant Size (mgd )
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 )
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% br 25 years. For I mgd plants, recalcination equipment is
8- 19
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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 — Chapter 8
1. CuIp, R. L., and CuIp, 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, 6110, 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 nigd 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/I is required for phosphate removal from typical
municipal wastewater while dosages of 50 to 100 mg/I 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 Al: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 f cc. Settling alone will reduce residual P to
about I mg/I.
If higher removals are required, filtration must be employed. With proper operation,
residual P may be reduced to less than 0.1 mg/I. Multimedia filters are preferred over
sand filters because of the extended length of run. Microstrainers have not been
successful in removing alum fcc.
Surface overflow rates for clarifiers have ranged from 580 to I ,440 gpd/ft 2 , 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/l ,000 gal.
depending on plant size and degree of treatment required.
9-I
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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 2 . Backwash with secondary
plant effluent was initiated either by headloss or by a rise in turbidity of the effluent.
Polyelectrolyte was used as a filter aid to control the depth of penetration of fioc into
the filter beds. With an alum dosage of 200 mg/I, 8 to 9 mg/I P in the secondary
plant effluent was reduced to 0.03 to 0.3 mg/I 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-I
1966 ESTIMATED COSTS FOR TERTIARY
PHOSPHATE REMOVAL WITH ALUM
Plant Capacity (nigd) Total Cost ( /l,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 polyelectrolyte are added. Sludge is recirculated into the coagulation zone to
improve efficiency of fcc 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/ft 2 . 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/I alum are shown in Table 9-2.
Table 9-2
NASSAU COUNTY PERFORMANCE DATA
Influent Effluent
Turbidity, JTU 20 - 25 0.2 - 1.5
Color, Pt-Co Units 20 - 40 none
COD, mg/I 78 13
P, mg/i 9 0.1 - 1.0
Costs have been estimated, for separate tertiary treatment only, at plant sizes of 1, 10,
and 100 mgd as shown in Table 9-3.
Table 9-3
NASSAU COUNTY COST ESTIMATES
Plant Capacity (mgd )
(1) (10) ( 100 )
(Costs, r// 1,000 gal)
Process
Coagulation 4.9 3.5 3.2
Filtration 1.8 1.1 1.0
Carbon adsorption 6.3 4.5 4.0
13.0 9.1 8.2
Operating labor 28.0 5.6 L8
Total Cost 41.0 14.7 10.0
9.3.3 Chicago, Hanover Park (9)
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 microstrutning for polishing. Average annual flow was I .5 rngd.
Combined cozigulalion plus sand Filtration removed 34, 58, and 82% P at dosages of
38. 60, .icid 140 mg/I of alum. respectively. Combined coagulation plus microstraining
elIecied .1 low pliu’ phate icmov l of 23% at an alum dosage of 38 mg/i. In general,
“I
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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 DaIlas(IO)
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/ft 2 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
4 gpm/ft . The biflow filter is designed for 100 gprn at a filtration rate of 8 gpm/ft 2 , or
4 gpmfft 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/I CaO 128 133
FeCI 3 Dose, mg/I Fe 3 25 56
Soluble Phosphate In, mg/i P 6.5 7.0
Soluble Phosphate Out, mg/I P 0.43 0.21
% Soluble P Removed 93.5 97.0
Test Period, days 12 2
9.3.5 Dayton(Il)
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 2 . Two filters designed for
loadings up to 8 gpm/ ft 2 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 I .80. Both air and water were used in backwashing.
Alum with activated silica or anionic polyelectrolyte as coagulant aids achieved
phosphate removals at a dosage rate of 5 lb alum/lb P0 4 3 down to residuals of 4
9-4
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mg/I P0 4 3 - in filter effluents (89% removal of phosphate). Further removals down to
I to 2 mg/I P0 4 3 required in excess of 200 mg/I alum. The clarifier was operated at
overflow rates of 655 to 1,035 gpd/ft 2 .
The filters removed 30 to 70% of the applied phosphate load. The dual media (liter
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 f 1,000 gal.
9.3.6 Oxidation Pond Effluents
At Lancaster, California, a research project was initiated to renovate wastewater for use
in recreational Iakes(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/I P with alum
dosages of 150, 200, and 300 mg/I, respectively, at an optimum p1 -I 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 0.1 5 mg/i p.
Final specifications for a full scale process included coagulation with about 300 mg/I
alum and 20 minutes of flocculation with a conventional paddle fiocculator,
sedimentation in a horizontal clarifier with an overflow rate of 400 gpd/t’t 2 , and
filtration through a dual media gravity filter consisting of IS 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 01
$ 184/I O 6 gal. The total cost of capital, operation, and maintenance or .s 3 mgd
facility was estimated at $ 150/106 gal.
Alum was recently evaluated by Shindala and Stewart(13) for phosphorus renioval and
polishing of waste stabilization pond effluents in Mississippi Jar tests were used in
evaluation. Using a minimum of 90% removal of phosphorus and 70% removal ol C ot)
as cnteria, the optimum alum dosage was 85 mg/I at pH 5.5. With higher doscs, 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 supernaiant
9.3.7 Moving Bed Filter (H)
Alum and polyelectrolyte were used to treat trickling filter effluent in a proprietary
moving bed filter at Liberty Corner, New Jersey. The device, a prodLict 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
floc-laden sand is removed mechanically at the top as required by pressure dicip.
9-5
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leaving a fresh filtration surface. The wastewater being tittered passes down through
the bed and flows through screens at the sides of the litter body.
A major factor controlling phosphate removal is the molar ratio of Al:?. At an alum
dosage of 200 mg I I and inuluent total phosphorus concentrations on the order of 25
to 28 mg/I as P (AI:P ratio = 0.6 to 0.7), the total phosphorus removal efliciency
averaged 90%. With lower total phosphorus concentrations (Al: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/I in the
effluent.
Capital costs for a 1 mgd moving bed filter are estimated at g264,000. Total operating
costs for such an installation using 200 mg/I alum are estimated at I2 /l,000 gal.
treated.
9-6
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9.4 References — Chapter 9
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
170 IOEKI, Contract 14-12-158, USD1, 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. CuIp, R. L., and Roderick, R. E., “The Lake Tahoe Water Reclamation Plant”,
JWPCF, 38:2, p 147 (1966).
6. CuIp, R. L., “Wastewater Reclamation at South Tahoe Public Utilities District”,
JAWWA, 60:1, p84(1968).
7. Rose, J. L., “Injection of Treated Waste Water into Aquifers”, Wat & Wastes Eng,
5:10, p40(1968).
8. Rose, J. L., “Advanced Waste Treatment in Nassau County, N. Y.”, Wat & Wastes
Eng., 7:2, p38(197 0 )
9. Lynam, B., Ettelt, G., and McAloon, T., “Tertiary Treatment at Metro Chicago by
Means of Rapid Sand Filtration and Microstramers”, 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)
II. 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 Stem, G., “Renovated Waste Water Creates Recreational Lake”,
Eni’ironmental 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 I7OIOEDO USD1,
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. Both of these chemicals are
available in either liquid or dry forms. Alum is acidic in nature while sodium
aluminate is alkaline and therefore one may have an advantage over the other
depending on the alkalinity of the wastewater.
10.1.1 Dry Alum
Properties and Availability. The commercial dry alum most often used in
wastewater treatment is known as “filter alum” and has the approximate chemical
formula Al 2 (S0 4 ) 3 . l4H 2 O and a molecular weight of 594. Alum is white to cream in
color and a 6% solution has a pH of about 3.4. The commercially available grades of
alum and their corresponding bulk densities and angles of repose are:
Grade Bulk Density (lb/ft 3 ) Angle of Repose
Lump 62 to 68
Ground 60 to 71 430
Rice 57 to 61 38°
Powdered 38 to 45 65°
Each of these grades has a minimum aluminum content of 17%, expressed as Al 2 0 3 .
Viscosity and solution crystallization 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 United States.
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 that the storage space is free of moisture.
Alum is shipped in 50 or 100 lb bags, drums, or in bulk by truck or rail. Bag
shipments may be ordered on wood pallets if desired. Locations of the major
production plants are listed in Table 10-1. At present, the price range for dry alum in
bulk quantities is $58 to $62/ton, F.O.B. the point of manufacture. Freight costs to
the point of usage must be added to this.
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Table 10-I
LOCATION OF ALUM MANUFACTURING PLANTS
Form of Alum
Location Manufacturer Available
ALABAMA
Coosa Pines American Cyanamid Liquid
Demopolis American Cyanamid Liquid
Mobile American Cyanamid Liquid and Dry
Naheola Stauffer Liquid
ARKANSAS
Pine Bluff Allied Liquid
CALIFORNIA
Bay Point (San Francisco) Allied Liquid and Dry
El Segundo (Los Angeles) Allied Liquid
Richmond (San Francisco) Stauffer Liquid
Vernon (Los Angeles) Stauffer Liquid
COLORADO
Denver Allied Liquid and Dry
FLORIDA
Fernandina Beach Tennessee Corp. Liquid
Jacksonville Allied Liquid
Port St. Joe Allied Liquid
GEORGIA
Atlanta (2 plants) Burns, Allied Liquid and Dry
Augusta Tennessee Corp. Liquid and Dry
Cedar Springs Tennessee Corp. Liquid
Macon Allied Liquid
Savannah Allied Liquid
ILLINOIS
E. St. Louis Allied Liquid and Dry
Joliet American Cyanamid Liquid and Dry
LOUISIANA
Bastrop Stauffer Liquid
Baton Rouge Stauffer Liquid
Monroe Allied Liquid
New Orleans Allied Liquid and Dry
Springhill Stauffer Liquid
MAINE
Searsport Northern Liquid and Dry
MARYLAND
Baltimore Olin Dry
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Table 10-I (Continued)
Form of Alum
Location Manufacturer Available
MASSACHUSETfS
Adams Holland Liquid
Salem Hamblet & Hayes Liquid and Dry
MICHIGAN
Detroit Allied Liquid
Escanaba American Cyanamid Liquid
Kalamazoo (2 plants) Allied, American Cyanamid Liquid
MINNESOTA
Cloquet American Cyanamid Liquid
Pine Bend North Star Liquid and Dry
MISSISSIPPI
Monticello American Cyanam id Liquid
Vicksburg Allied Liquid
NEW JERSEY
Newark Essex Liquid and Dry
Warners American Cyanamid Liquid and Dry
NORTH CAROLINA
Acme Wright Liquid
Plymouth American Cyanamid Liquid
OHIO
Chillicothe Allied Liquid
Cleveland Allied Liquid and Dry
Hamilton American Cyanamid Liquid and Dry
Middletown Allied Liquid
OREGON
North Portland Stauffer Liquid and Dry
PENNSYLVANIA
Johnsonburg Allied Liquid
Marcus Hook Allied Liquid and Dry
Newell Allied Liquid
SOUTH CAROLINA
Catawba Burns Liquid
Georgetown American Cyanamid Liquid and Dry
TENNESSEE
Chattanooga American Cyanamid Liquid and Dry
Counce Stauffer Liquid
Springfield Burns Liquid
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Table 10-I (Continued)
Location
TEXAS
Manufacturer
Form of Alum
Available
Houston (2 plants)
VIRGINIA
Covi ngton
Hopewell
Norfolk
WASHINGTON
Kennewick
Tacoma (2 plants)
Vancouver
WISCONSIN
Men a sha
Wisconsin Rapids
Stauffer, Ethyl
Allied
Allied
1 -lowerton Gowen
Allied
Stauffer, Allied
Allied
Allied
Allied
Liquid and Dry
Liquid
Liquid
Liquid
Liquid
Liquid
Liquid and Dry
Liquid
Liquid
Manufacturers and Addresses
Allied Chemical Corporation
Industrial Chemicals Division
P. 0. Box 1139R
Morristown, New Jersey 07960
American Cyanamid Company
lnd. Chem. Div.
P. 0. Box 66189
Chicago. Illinois 60666
Burns Chemical Company
Charleston, South Carolina
Essex Chemical Corporation
1402 Broad Street
Clifton, Ncw Jersey 07015
Ethyl Corporation
Houston. Tcxas
Uamb et & Hayes Company
Colonial Road
Salem, Massachusetts 10970
HoIJ.mcl Chemical Company
Adams, Massachusetts
Howerton Gowen Company, Inc.
Norfolk, Virginia
Northern Chemical Industries, Inc.
Searsport, Maine
North Star Chemicals Inc.
P 0. Box 28-T
South St. Paul, Minnesota
Olin Corporation,
Chemicals Division
745 Fifth Avenue
New York, New York
Stauffer Chemical Company
299 Park Avenue
New York, New York 10017
Tennessee Corporation
Cities Service Company
Industrial Chemicals Division
P. 0. Box 50360
Atlanta, Georgia
Wright Chemical Co.
Acme, North Carolina
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General Design Considerations. Ground and rice alum are the grades most
commonly used by utilities 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 detrimental to feeding because of the possibility of packing the bin.
Alum dust is present in the ground grade and will cause minor irritation of the eyes
and nose on breathing. A respirator may be worn for protection against alum dust.
Gloves may be worn to protect the hands. Because of minor irntation in handling and
the possibility of alum dust causing rusting of adjacent machinery, dust removal
equipment is desirable. Alum dust should be thoroughly flushed from the eyes
immediately and washed from the skin with water.
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. 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 as the alum
can contain abrasive impurities.
Bags and drums of alum should be stored in a dry location to avoid caking. Bag or
drum loaded hoppers should have a nominal storage capacity for eight hours at the
nominal maximum feed rate so that personnel are not required to charge the hopper
more than once per shift. Converging hopper sections should have a minimum slope of
60° to prevent arching.
Bulk storage hoppers should terminate at a bin gate so that the 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.
Feeding Equipment. The feed system includes all of the components required
for the proper 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 initial low cost and usually lower capacities
are the basis of selection. Volumetric feeder mechanisms are usually exposed to the
corrosive dissolving chamber vapors which can cause corrosion of discharge areas.
Manufacturers usually control this problem by use of an electric heater to keep (lie
feeder housing dry or by using plastic components in the exposed areas.
Volumetric dry feeders presently in general use are 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 hopper 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
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PIPE (PNEUMATIC)
FIGURE 10-I TYPICAL DRY FEED SYSTEM
GRAVITY TO
APPLICATION
PUMP
TO APPLICATION
BULK STORAGE
BIN
DUST COLLECTOR
DAY HOPPER
FOR DRY CHEMICAL
FROM BAGS OR
BAG FILL
.SCREEN
WITH BREAKER
.BIN GATE
FLEX I BLE
CONNECTION
DUST
ALTERNATE SUPPLIES DEPENDING
SCALE OR SAMPLE CHUTE
WATER SUPPLY
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opposite directions so that the turning and reciprocating motion alternately fills one
half of the screw while the other half of the screw 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 screw. The accuracy of volumetric feeders
is influenced by the character of the material being fed and ranges between ±1% for
free-flowing materials and ±7% for cohesive matenals. This accuracy is volumetric and
should not be related to accuracy by weight (gravimetnc).
Where the greatest accuracy and the most economical use of chemical is desired, the
loss-in-weight type feeder should be selected. This feeder is limited to the low and
intermediate feed rates up to a maximum rate of approximately 4,000 lb/hour. The
loss-in-weight type feeder consists of a matenal hopper and feeding mechanism
mounted on enclosed scales. The feed rate controller retracts the 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 the output of the feeding mechanism Continuous comparison of
actual hopper weight with set hopper weight prevents cumulative errors. Accuracy of
the loss-in-weight feeder is ±1% by weight of the set rate.
Belt type gravimetric feeders span the capacity ranges of volumetric and loss-in-weight
feeders and can usually be sized for all applications encountered in wastewater
treatment applications. Initial expense is greater than for the volumetric feeder and
slightly less than for 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, or the belt speed, or both. Controllers
in general use are mechanical, pneumatic, electric and mechanical-vibrating. Accuracy
specified for belt type gravimetnc feeders should be within ±1% of set rate. Materials
of construction of feed equipment normally include 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), or plastics.
Dissolvers should be sized for preparation of the desired solution strength. The solution
strength usually recommended is 0.5 lb of alum to I gal. 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 that the water to alum ratio may be properly established and controlled.
Piping and Accessories. FRP, plastics (polyvinyl chloride, polyethylene,
polypropylene, and other similar materials), and rubber are in general use and are
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 that
objectionable deflection will not be expenenced. The alum solution should be injected
into a zone of rapid mixing or turbulent flow.
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Solution flow by gravity to the point of discharge is desirable. When gravity flow is
not possible, transfer components should be selected that 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 further in the section
on liquid alum.
Valves used in solution lines should be plastic, type 31 6 stainless steel or rubber-lined
iron or steel.
Pacing and Control. Standard instrument control and pacing signals are generally
acceptable for common feeder system operation. Volumetric and gravimetric feeders are
usually adaptable to operation from any standard instrument signals.
When solution must be pumped, consideration should be given to use of holding tanks
between the dry feed system and feed pumps, and the 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 then transfer the alum
stock solution to the point of application without further dilution.
Means should be provided for calibration of the chemical feeders. Volumetric feeders
may be mounted on platform scales. Belt feeders should include a sample chute and
box to catch samples for checking 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
Properties and Availability. Liquid alum is shipped in rubber lined or stainless
steel, insulated tank cars or trucks. Alum shipped during the winter is heated prior to
shipment so that crystallization will not occur during transit. Liquid Alum is shipped
at a solution strength of about 8.3% as A1 2 0 3 or about 49% as A1 2 (S0 4 ) 3 . 14H 2 0.
This solution has a density of 11.1 lb/gal. at 60uF and contains about 5.4 lb dry alum
(17% Al 2 0 3 ) per gal. of liquid. This solution will begin to crystallize at 18°F.
Crystallization temperatures of other solution strengths are as follows (1):
Temperature of
% Al O 3 Crystallization, ° F.
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
10-8
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The viscosity of various alum solutions is given in Figure 10-2.
Tank truck lots of 3,000 to 5,000 gal. are available. Tank car lots are available in
quantities of 7,000 to 18,000 gal. Production locations of liquid alum are listed in
Table 10-1. The current price range of liquid alum on an equivalent dry alum (17%
Al 2 0 3 ) basis is about $43 to $48/ton, F.O.B. the point of manufacture. Liquid alum
will generally be more economical than dry alum if the point of use is within a 50 to
100 mile radius of the manufacturing plant.
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.
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 180 F
to prevent crystallization. Storage tanks should be constructed of type 3 16 stainless steel;
FRP; steel lined with rubber, polyvinyl chloride, or lead. 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’/2 times the quantity
of shipments. A 10 day to 2 week supply should be provided to allow for unforeseen
shipping delays.
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 be supplied by a centrifugal pump. The latter arrangement requires an excess
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FIGURE 10-2 VISCOSITY OF ALUM SOLUTIONS
(Courtesy of Al lied Chemical Co.)
10
8.0
6.0
‘hO
(I,
LU
U i
C.,
>-
I—
C ,)
0
C.)
C l)
TEMPERATURE, °F
10 - 10
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FIGURE 10-3 ALTERNATIVE LIQUID FEED SYSTEMS
FOR OVERHEAD STORAGE
GRAVITY FEED
PRESSURE FEED
FIGURE 10- I l
ALTERNATIVE
FOR GROUND
FLOAT
VALVE
LU
LU
LU
ROTODIP—TYPE FEEDER
GRAVITY FEED GRAVITY FEED
TAMETER ‘METERING PUMP
GRAVITY FEED
TRANSFER
LU
-J
LU
LU
!
LIQUID FEED SYSTEMS
STORAGE
10 - 11
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flow return line to the storage tank, as shown by Figure 104. Centrifugal pumps
should be direct-connected but not close-coupled because of possible leakage into the
motor, and should be constructed f 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 stamless steel, FRP, plastics, and rubber.
Piping and Accessories. Piping systems for alum should be FRP, plastics (subject
to temperature limits), type 31 6 stainless steel, or lead. Piping and valves used for
alum solutions are also discussed in the preceding section on dry alum.
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
Properties and Availability. Dry sodium aluminate, Na 2 A 1 2 O 4 , is available from
two major manufacturers in the United States. They are:
Nalco Chemical Company
Chicago, Illinois (Plant located at Clearing, Ill.)
Reynolds Chemicals
Richmond, Virginia (Plant located at Bauxite, Ark.)
Sodium aluminate is shipped in 50 lb bags and has a bulk density ranging from 40 to
50 lb/ft 3 . The A 1 2 0 3 content ranges from 41 to 46%. 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 current price of dry sodium aluminate in 40,000 lb shipments ranges from $0.1 I
to $0.14/Ib, F.O.B. the point of manufacture. Prices increase substantially for smaller
shipments.
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.
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Contact with skin, eyes, and clothing should be avoided. Aluminate dust or solution
spray should not be breathed.
Storage. Dry sodium aluminate is stored as received, in bags, and at optimum
conditions of 60 to 900 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 detenorates on exposure to the atmosphere and
care should be taken to avoid tearing of bags.
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 lb/gal. or 6% solution strength with a dissolver detention time of 5
minutes at the maximum feed rate.
After dissolving 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.
Piping and Accessories. Materials for piping and transporting 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.
Pacing and Control. Pacing and control fundamentals are similar to those
described for dry alum.
The amount of dilution does not appear to be a consideration in the use of sodium
aluminate. Therefore, the use of float valves to satisfy centrifugal pump suction, and
the use of eductors are permissible.
10.1.4 Liquid Sodium Aluminate
Properties and Availability. Liquid sodium aluminate is available from the
following major manufacturers m the United States:
Nalco Chemical Company
Chicago, Illinois
Vinings Chemical Company
Vinings, Georgia
Conservation Chemical Company
Kansas City, Missouri
There is considerable variety in the composition of sodium aluminate from the
manufacturers listed. The A1 2 0 3 content varies from 4.9 to 26.7%. The lower solution
10- 13
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strengths are usually more expensive because of the cost of freightmg 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 gal. drums (380 Ib), tank truck, and tank
car quantities. The current price of sodium aluminate in 40,000 lb quantities is about
$0.08/lb (A1 2 0 3 content, 26.4%) F.O.B. the point of manufacture.
General Design Considerations. Matenal selection and dilution restrictions are
not as limited as for liquid alum, because of the noncorrosive nature of sodium
alummate. Sodium aluminate is a strong alkali and the same precautions should be
exercised in handling it as in handling caustic soda.
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
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 the CO 2 . Steam injection facilities are required at the unloading site.
Feeding Equipment. Feeding equipment and systems as described for liquid
alum generally apply to sodium aluminate except with changes of requirements
regarding dilution and materials of construction as described above.
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.
Piping and Accessories. Material requirements are the same as previously
indicated for solutions of dry aluminate.
Pacing and Control. System pacing and control requirements are the same as
described for liquid alum.
10.1 .5 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% A1 2 0 3 ) for
phosphorus precipitation. A mole ratio of Al:P of 2.0:1 was assumed to treat an
influent phosphorus concentration of 10 mg/I as P. This corresponds to an alum dose
of 191 mg/I (17.4 mg/i as AI 3 ).
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 source of chemical supply. Piping and buildings to house the
10- 14
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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:
Plant Size Estimated Cost
(mgd) ($)
15,000
10 47,500
100 588,000
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 gal. FRP
tanks with accessories are provided for storage of liquid alum. The total capacity of
6,000 gal. allows some flexibility in ordering and receiving 4,000 gal. tank truck
shipments.
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 gal. 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 gal. underground concrete tanks with a rubber lining. The
underground storage necessitates transfer pumps and day tanks for the rotodip feeders
but eliminates the need for heating the tanks. Three 500 gal. 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
Properties and Availability. Liquid ferric chloride is a corrosive, dark brown
10- 15
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oily-appearing solution having a density as shipped and stored of 11.2 to 12.4 lb/gal.
(35 to 45% FeCI 3 ) (2). The ferric chloride content of these solutions, as FeCl 3 , is
3.95 to 5.58 lb/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 1 62.22. Viscosities of ferric chloride
solutions at various temperatures are presented m Figure 10-6.
Liquid fernc chloride is shipped in 3,000 to 4,000 gal. bulk truckload lots, in 4,000
to 10,000 gal. bulk carload lots, and in 5 and 13 gal. carboys. Liquid ferric chloride is
produced at the following locations.
Dow Chemical Co.
Midland, Michigan
Pennwalt Corp.
Philadelphia, Pa. (Plant at Wyandotte, Mich.)
The current price of liquid ferric chloride in bulk quantities is about $0.04 to
$0.045/lb (as Fed 3 ), F.O.B. the point of manufacture.
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.
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 fernc chloride is mgested, induce
vomiting and call a physician.
Storage. Ferric chlonde 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. Fernc chloride
can be stored for long periods of time without deterioration. The total storage
capacity should be lY2 times the largest anticipated shipment, and should provide at
least a 10 day to 2 week supply of the chemical at the design average dosage.
10- 16
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I I I I I I I
A Agitated Solutions may begin to develop crystals —
below this line —
B Unagitated Solutions begin to develop crystals when
the bulk solution temperatqre drops to about this line.
Ice crystals form Delow 33% FeC1 3 & FeCl 6H 2 0 —
Crystals form above 33% FeCI 3 .
A
,
,
,
8/ -
/
/
/
/
/
/
/
/
/
I —
I
I —
I
I —
If I I I I I I I
314 38 ‘42 146 50
FIGURE 10-5 FREEZING POINT CURVES FOR COMMERCIAL
FERRIC CHLORIDE SOLUTIONS
(Courtesy of Dow Chemical Co.)
A
60
50
‘40
30
LI
o 20
‘ -10
-20
-30
- ‘ I D
—50
I I I I I
10 1 14 18 22 26 30
% FeC1 3
10-17
-------
LU
a
I-
I-
LU
C-)
>-
I-
C ,,
a
C-,
C l)
C-,
I-
% FeC1 3
FIGURE 10-6 VISCOSITY VS COMPOSITION OF FERRIC CHLORIDE
SOLUTIONS AT VARIOUS TEMPERATURES
(Courtesy of Dow Chemical Co.)
0 10 20
30
50
10 - 18
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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 fernc 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 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.
Piping and Accessories. Materials for piping and transporting ferric chloride
should be rubber or Saran lined steel, hard rubber, FRP, or plastics. Valving 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.
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)
Properties and Availability. Ferrous chloride, FeC I 2 , as a liquid is available in
the form of waste pickle liquor from steel processing. The liquor weighs between 9.9
and 10.4 lb/gal. and contains 20 to 25% Fed 2 or about 10% available Fe 2 . A 22%
solution of Fed 2 will crystallize at a temperature of .40 F. The molecular weight of
Fed 2 is 126.76. Free acid in waste pickle liquor can vary from I to 10% and usually
averages about 1 .5 to 2.0%. Ferrous chloride is slightly less corrosive than ferric
chloride.
Waste pickle liquor is available in 4,000 gal. 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 an FeCl 2 content of about 22% at a price of $0.04/lb of FeC ! 2 in bulk
car or truckload quantities, F.O.B. Midland, Michigan.
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.
10.2.3 Ferric Sulfate
Properties and Availability. Ferric sulfate is marketed as a dry, partially
10- 19
-------
hydrated product with a maximum of three water molecules. Density, % Fe 2 (SO 4 ) 3 ,
and molecular weight of the two most common hydrates are presented below:
Properties of Two Ferric Sulfate Hydrates (3)
Fe 2 (S0 4 ) 3 .3 H 2 0 Fe 2 (S0 4 ) 3 .2 H 2 0
Density 63 to 72 lb/ft 3 83 to 97 lb/ft 3
% Fe 2 (SO 4 ) 3 68% 76%
Molecular weight 453.8 435.8
Fernc sulfate is shipped in car and truck load lots of 50 lb and 100 lb moistureproof
paper bags and 400 lb fiber drums. Bulk carload shipments in box and closed hopper
cars are available. Major producers and corresponding plant locations are:
Producer Plant Location
Stauffer Chemical Co. San Francisco, California
Fort Worth, Texas
Tennessee Corp. (Cities Service Co.) Copper Hill, Tennessee
The current price of ferric sulfate (21.8% Fe) is about $38/ton, F.O.B. Copper Hill,
Tennessee.
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.
General Design Considerations. Aeration of ferric sulfate should be held to a
minimum because of the hygroscopic nature of the material, particularly m 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 m 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.
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.
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.
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
10 - 20
-------
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.
Piping and Accessories. Piping systems for fernc sulfate should be FRP, plastics,
type 316 stainless steel, rubber or glass.
Pacing and Control. System pacing and control are the same as discussed for
dry alum.
10.2.4 Ferrous Sulfate
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 FeSO 4 .7H 2 0, with a molecular weight
of 278, and containing 55 to 58% FeSO 4 and 20 to 21% Fe. The product has a bulk
density of 62 to 66 lb/ft 3 . 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 Ib) and bag (50 and 100 Ib) shipments are available from producers
at the following locations:
American Cyanamid Co. Savannah, Georgia
Byproducts Processing Ca, Inc. Baltimore, Maryland
Glidden Co. Baltimore, Maryland
Cosmin Corp. Baltimore, Maryland
National Lead St. Louis, Missouri
National Lead Sayreville, New Jersey
The current price of ferrous sulfate in bulk carload and truckload quantities is about
$ 18/ton (21% Fe).
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 680 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 - 21
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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 lb/gal. or 6%
with a detention time of 5 minutes in the dissolver. 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 Femc Chloride
Feed Systems to Existing Wastewater Treatment Plants
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% Fed 3 ) for
phosphorus precipitation. A mole ratio of Fe:P of 2.0: 1 was assumed to treat an
influent phosphorus concentration of 10 mg/i as P. This corresponds to a ferric
chloride dose of 104 mg/i (36 mg/i as Fe 3 ).
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 source of chemical supply. Pipmg 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:
Plant Size Estimated Cost
(mgd) ($)
15,000
10 45,000
100 494,000
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 gal. FRP tanks with accessories. The 6,000 gal. total
capacity allows use of liquid ferric chloride in 4,000 gal. 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.
10 - 22
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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 gal. rubber-lined concrete tanks located underground. The underground storage
necessitates transfer pumps and day tanks for the rotodip feeders, but eliminates the need
for heating the tanks. Three 500 gal. 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 nigd 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
Properties and Availability. Quicklime, CaO, has a density range of approximately
55 to 70 lb/ft 3 , 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 truck load lots. It is also shipped in 80
and 100 lb multiwall “moistureproof” paper bags. Lime is produced at the locations
indicated by Table 10-2.
Table 10-2
LOCAT 1ON OF LIME MANUFACTURING PLANTS (6)
Form of
Location Manufacturer Lime Available
ALABAMA
A llgood Cheney Lime & Cement Co. High Calcium
Keystone Southern Cement Co., High Calcium
Div. Martin Marietta Corp.
Landmark Cheney Lime & Cement Co. High Calcium
Montevallo U. S. Gypsum Co. High Calcium
Roberta Southern Cement Co. High Calcium
Div. Martin Marietta Corp.
Saginaw Longview Lime Co., Div. High Calcium
Woodward Co., Div. Mead Corp.
Siluria Alabaster Lime Co. High Calcium
ARIZONA
Douglas Paul Lime Plant, inc. High Calcium
Globe Hoopes & Co. High Calcium
Nelson U. S. Lime Div., The Flintkote Co. High Calcium
10 - 23
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Table 10-2 (Continued)
Form of
Location Manufacturer Lime Available
ARKANSAS
Batesville Batesville White Lime Co., High Calcium
Div. Rangaire Corp.
CALIFORNIA
City of Industry U. S. Lime Div., The Flintkote Co. High Calcium
Diamond Springs Diamond Springs Lime Co. High Calcium
Lucerne Valley Pfizer, Inc., Minerals, Pigments High Calcium
and Metals Div.
Richmond U. S. Lime Div., The Flintkote Co. High Calcium
Salinas Kaiser Aluminum & Chemical Corp. Dolomitic
(currently captive lime)
Sonora U. S. Lime Div., The Flintkote Co. Dolomitic
Westend Stauffer Chemical Co. High Calcium
COLORADO
Ft. Morgan Great Western Sugar Co. High Calcium
CONNECTICUT
Canaan Pfizer, Inc., Minerals, Pigments Dolomitic
and Metals Div.
FLORIDA
Brooksville Chemical Lime Co. High Calcium
Sumterville Dixie Lime and Stone Co. High Calcium
ILLINOIS
Marblehead Marblehead Lime Co. High Calcium
McCook Standard Lime & Refractories Dolomitic
Div., Martin Marietta Corp.
Quincy Marblehead Lime Co. High Calcium
So. Chicago Marblehead Lime Co. High Calcium
Thornton Marblehead Lime Co. Dolomitic
INDIANA
Buffington Marblehead Lime Co. High Calcium
IOWA
Davenport Linwood Stone Products Co., Inc. High Calcium
KENTUCKY
Carntown Black River Mining Co. High Calcium
LOUISiANA
Morgan City Pelican State Lime Corp. High Calcium
New Orleans U. S. Gypsum Co. High Calcium
10 - 24
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Table 10-2 (Continued)
Form of
Location Manufacturer Lime Available
MARYLAND
LeGore LeGore Lime Co. High Calcium
Woodsboro S. W. Barrick & Sons, Inc. High Calcium
MASSACHUSEflS
Adams Pfizer, Inc., Minerals, Pigments High Calcium
and Metals Div.
Lee Lee Lime Corp. Do lomitic
MICHIGAN
Detroit Detroit Lime Co. High Calcium
Ludington Dow Chemical Co. (currently captive lime) High Calcium
Menominee Limestone Products Co., Div. High Calcium
C. Reiss Coal Co.
River Rouge Marblehead Lime Co. High Calcium
MINNESOTA
Duluth Cutler Magner Co. Nigh Calcium
MISSOURI
Bonne Terre Valley Dolomite Co. Dolomitic
Hannibal Marblehead Lime Co. High Calcium
Ste. Genevieve Mississippi Lime Co. High Calcium
Springfield Ash Grove Cement Co. High Calcium
NEVADA
Apex Li. S. Lime Div., The Flintkote Co. High Calcium
Henderson U. S. Lime Div., The Flintkote Co. Dolomitic &
High Calcium
McGil l Morrison-Weatherly Corp. High Calcium
Sloan U. S. Lime Div., The Flintkote Co. Dolomitic &
High Calcium
NEW JERSEY
Newton Limestone Products Corp. of America High Calcium
OHIO
Ash tabula Union Carbide Olefins Co. High Calcium
Carey National Lime & Stone Co. Dolomitic
Cleveland Cuyahoga Lime Co. High Calcium
Delaware Marble Cliff Quarries Co. High Calcium
Geona U. S. Gypsum Co. Dolomitic
Gibsonburg (2 plants) Pfizer, Inc., Mjnerals, Pigments Dolomi tic
and Metal Div., National Gypsum Co.
Huron Huron Lime Co. High Calcium
10 - 25
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Table 10-2 (Continued)
Form of
Location Manufacturer Lime Available
OHIO (Continued)
Marble Cliff Marble Cliff Quarries Co. High Calcium
Millersville J. E. Baker Co. Dolomitic
Woodville Ohio Lime Co., Standard Lime & Re- Doloniitic
fractories Div. , Martin Marietta Corp.
OKLAHOMA
Marble City St. Clair Lime Co. High Calcium
Sallisaw St. Clair Lime Co. High Calcium
OREGON
Baker Chemical Lime Co. of Oregon High Calcium
Portland Ash Grove Cement Co. High Calcium
PENNSYL VAN IA
Annville Bethlehem Mines Corp. High Calcium
Bellefonte (2 plants) National Gypsum Co., Warner Co. High Calcium
Branchton Mercer Lime & Stone Co. High Calcium
Devault Warner Co. Dolomitic
Everett New Enterprise Stone & Lime Co. High Calcium
Pleasant Gap Standard Lime & Refractories Div ., High Calcium
Martin Marietta Corp.
flymouth Meeting G. & W. H. Corson, Inc. Do lomitic
SOUTH DAKOTA
Rapid City Pete Lien & Sons, inc. High Calcium
TENNESSEE
Knoxville (2 plants) Foote Mineral Co., Williams Lime High Calcium
Manufacturing Co.
TEXAS
Blum Round Rock Lime Companies High Calcium
Cleburne Texas Lime Co., Div. Rarigaire Corp. High Calcium
Clifton Clifstone Lime Co. High Calcium
Houston U. S. Gypsum Co. High Calcium
McNe i l Austin White Lime Co. High Calcium
New Braunfels U. S. Gypsum Co. High Calcium
Round Rock Round Rock Lime Companies High Calcium
San Antonio McDonough Bros., Inc. High Calcium
UTAH
Grantsville U. S. Lime Div., The Ftiritkote Co. Dolomitic &
High Calcium
Lehi Rollins Mining Supplies Co. High Calcium
10 - 26
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Table 10-2 (Continued)
Form of
Location Manufacturer Lime Available
VERMONT
Winooski Vermont Assoc. Lime Industries, Inc. High Calcium
VIRGINIA
Clearbrook W. S. Frey Co., Inc. High Calcium
Kimballton (2 plants) Foote Mineral Co., National Gypsum High Calcium
Company
Stephens City M. J. Grove Lime Co., Div. The High Calcium
Flintkote Co.
Strasburg Chemstone Corp. High Calcium
WASHINGTON
Tacoma Domtar Chemicals Inc. High Calcium
WEST VIRGINIA
Miliville Standard Lime & Refractories Div., Doloniltic
Martin Marietta Corp.
Riverton Germany Valley Limestone Div., High Calcium
Greer Steel Co.
WISCONSIN
Eden Western Lime & Cement Co. Dolomitic
Green Bay Western Lime & Cement Co. Uigh Calcium
Knowles Western Lime & Cement Co. Dolornitic
Manitowoc Rockwell Lime Co. Dolomitic
Superior Cutler-LaLiberte-M cDougall Corp. High Calcium
Current prices for bulk pebble quicklime range from $ 17/ton to $2 i/ton with the higher
prices generally in the far west, and higher than average in the north. Bagging adds
approximately $4/ton to the cost.
The CaO content of commercially available quicklime can vary quite widely over an
approximate range of 70 to 96%. Content below 88% is generally considered below
standard in the municipal use field (5). 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 - 27
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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. I-topper agitation 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.
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.
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.
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 maintam 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 evolves heat in hydrating the CaO to Ca(OH) 2 , and
therefore, vapor removers are required for feeder protection.
10 - 28
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DUST CO1LECTOR
NOTE: VAPOR REMOVER
NOT SHOWN FOR CLARITY
SCALE
OR SN4PLE
SOLENOID
FILL PIPE (PNEUMATIC)
N GATE
FLOW RECORDER
WITh PACING
pH RECORDER
CONTROLLER
-
PRESSURE
VALVE FEED
PRESSURE
VALVE
FIGURE 10-7 TYPICAL LIME FEED SYSTEM
GRAVITY FEED
RECI R ULATI ON
PUMP
TANK
PU’,
tO 29
-------
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 pipmg 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.
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
p1-I 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 1 I .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 1 0-7 illustrates such a control system.
10.3.2 Hydrated Lime
Properties and Availability. Hydrated lime, Ca(OH) 2 , has a working density of
25 to 35 lb/ft 3 , 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 cost of bulk hydrated lime varies from $18 to $22/ton. Bagged lime is
available but increases the cost about $4/ton. The availability of hydrated lime may be
determined by contacting manufacturers listed in Table 10-2. The pH of a saturated,
hydrated lime solution is the same as that given for quicklime.
General Design Considerations. Hydrated lime is slaked lime and needs only
enough water added to form milk of lime. Wetting or dissolving chambers are
10 - 30
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usually designed to provide 5 minutes detention with a ratio of 0.5 lb per gal. of
water or 6% slurry at the maximum feed rate. Hydrated lime is usually used where
maximum feed rates do not exceed 250 lb/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.
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.
Feed Equipment. Volumetnc 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 m 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.
Piping and Accessories. Pipmg and accessories as described for quicklime are
also appropriate for hydrated lime.
Pacing and Controls. 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.
10.3.3 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 1 50 mg/I 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 mcluded 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 for engineering and
contingencies. The estimated costs for each size plant are:
10-31
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Additional Cost of
Estimated Flow and pH Total
Plant Size Base Bid Controls Cost
(mgd) ($) (S) ($)
15,800 3,500 19,300
10 80,100 13,200 93,300
100. 332,200 92,700 424,900
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 truck loads of quicklime each, and it is assumed that the truck will be equipped
with a pneumatic blower. Equipment for the 100.0 mgd plant includes four gravimetric
feeder-slaker systems. The base estimate includes the bin gate and feeder-slaker
accessories and controls.
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 Other Compounds for pH Adjustment
10.4.1 Soda Ash
Properties and Availability. Soda ash, Na 2 CO 3 , is available in two forms. Light
soda ash has a bulk density range of 35 to 50 lb/ft 3 and a working density of 41
lb/ft 3 . Dense soda ash has a density range of 60 to 76 lb/ft 3 and a working density
of 63 lb/ft 3 . The pH of a 1% solution of soda ash is 11.2.
The molecular weight of soda ash is 106. Commercial punty ranges from 98 to 99%
Na 2 CO 3 . 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 m bulk by truck, box car and hopper car, and in 100 lb bags
from the following locations:
Location Manufacturer
CALIFORNIA
Bartlett PPG industries, Inc.
Trona American Potash and Chemical Corp.
Westend Stauffer Chemical Co.
GEORGIA
Brunswick Allied Chemical Co.
LOUISIANA
Baton Rouge Allied Chemical Co.
Lake Charles Olin Chemicals
10 - 32
-------
a 2 CO 3
FIGURE 10-8 VISCOSITY OF SODA ASH SOLUTIONS
(Courtesy PPG Industries Inc., Chemical Div.)
C ,,
LU
C , ,
o
0
LU
I-
C,,
0
C.)
C,)
0 5 10 15 20 25 30
- 33
-------
Location Manufacturer
MICHIGAN
Wyandotte Wyandotte Chemicals Corp.
NEW YORK
Solvay Allied Chemical Co.
OHIO
Barberton PPG Industries, Inc.
Painesville Diamond Shamrock Chemical Co.
TEXAS
Corpus Christi PPG Industries, Inc.
WEST VIRGINIA
Moundsvile Allied Chemical Co.
WYOMING
Green River (3 plants) Allied Chemical Co, FMC Corp., and
Stauffer Chemical Corp.
The current price for soda ash ranges from $40 to $45/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.
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
lb of dense soda ash per gal. 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.
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 CO 2 and water forming the less active sodium
bicarbonate (NaHCO 3 ). Material recommended for unloading facilities is steel.
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.
Piping and Accessories. Materials of construction for piping and accessories
should be iron, steel, rubber, and plastics.
Pacing and Control. Controls as discussed for dry alum apply also to soda ash
equipment.
10 - 34
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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.
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 lb/gal. for the
50% solution and 14.18 lb/gal. for the 73% solution. These solutions contain 6.38 lb/gal.
NaOH and 10.34 lb/gal. NaOH, respectively. The crystallization temperature is 53° F for
the 50% solution and I 65° 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.
Truck load lots of 1,000 to 4,000 gal. are available in the 50% concentration only.
Both shipping concentrations can be obtained in 8,000, 10,000, and 16,000 gal. car
load 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 Table
10-3. The current price for liquid caustic soda ranges from about $65 to $70/ton
(NaOH), F.O.B. the point of manufacture.
Table 10-3
LOCATION OF CAUSTIC SODA MANUFACTURING PLANTS
Location Manufacturer
ALABAMA
Lemoyne (Mobile) Stauffer
Mcintosh Olin
Muscle Shoals Diamond Shamrock
CALIFORNIA
Pittsburg Dow
DELAWARE
Delaware City Diamond Shamrock
GEORGIA
Augusta Olin
Brunswick Allied
KANSAS
Wichita Vulcan
KENTUCKY
Calvert City (2 plants) Pennwalt, Goodrich
10 - 35
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0.
0
0.
0
FIGURE 10-9 VISCOSITY OF CAUSTIC SODA SOLUTIONS
(Courtesy of Hooker Chemi.cal Co.)
U)
w
U)
0
I-
uJ
I-
U)
0
U)
70 80 90 100 110 120 130 1UO 150 160 170 180 190 200
TEMPERATURE, °F
10 - 36
-------
Table 10-3 (Continued)
Location Manufacturer
LOUISIANA
Baton Rouge Allied
Geismar Wyandotte
Lake Charles (2 plants) PPG, Olin
Plaquemine Dow
St. Gabriel Stauffer
Taft Hooker
MICHIGAN
Midland Dow
Montague Hooker
Wyandotte (2 plants) Pennwalt, Wyandotte
NEVADA
Henderson Stauffer
NEW JERSEY
Linden GAF
NEW YORK
Niagara Falls (3 plants) Hooker, Olin, Stauffer
Solvay Allied
NORTH CAROLINA
Acme Allied
OHIO
Barberton PPG
Cleveland Harshaw
Painesville Diamond Shamrock
OREGON
Portland Pennwalt
TENNESSEE
Charleston Olin
TEXAS
Corpus Christi PPG
Deer Park (Houston) Diamond Shamrock
Freeport Dow
Port Neches Jefferson
VIRGINIA
Saltville Olin
WASHINGTON
Tacoma (2 plants) Hooker, Pennwalt
10 - 37
-------
Table 10-3 (Continued)
Location
WEST VIRGINIA
Moundsvile
Natnum
South Charleston
Manufacturers and addresses
Manufacturer
Allied
PPG
FMC
Allied Chemical
Solvay Process Division
40 Rector Street
New York, New York
Diamond Shamrock Chemical Co.
300 Union Commerce Building
Cleveland, Ohio 44115
Dow Chemical Co.
Abbott Road
Midland, Michigan 48640
GAF Corp. Chemical Division
140 West 5 I st Street
New York, New York
FMC Corporation
Inorganic Chemicals Div.
633 Third Avenue
New York, New York 10017
B. F. Goodrich Chemical Co.
3135 Euclid Avenue
Cleveland, Ohio
Harshaw Chemical Co.
1945 East 97th Street
Cleveland, Ohio 44106
Hooker Chemical Corp.
P. 0. Box 344
Niagara Falls, New York 14302
Jefferson Chemical Co., Inc.
3336 Richmond Avenue
Houston, Texas 77006
Olin Corporation
Chemicals Division
745 Fifth Avenue
New York, New York
Pennwalt Corporation
Pennwalt Building
Three Penn Center
Philadelphia, Pa. 19102
PPG industries, Inc.
Chemical Division
I Gateway Center
Pittsburgh, Pa. 15222
Stauffer Chemical Co.
Industrial Chemical Div.
299 Park Avenue
New York, New York
Vulcan Materials Co.
Chemicals Division
P. 0. Box 545-T
Wichita, Kansas 67201
Wyandotte Chemicals Corp.
Michigan Alkali Division
Wyandotte, Michigan
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
10 - 38
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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
possibie. Other parts of the body should be protected as necessary to prevent alkali
burns. Areas exposed to caustic soda should be washed with copius 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.
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.
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.
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):
10 - 39
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K FILL LINE
VENT, OVERFLOW
AND DRAIN
DIUM HYDROXIDE
FEEDER
FIGURE 10-10 TYPICAL CAUSTIC SODA FEED SYSTEM
STORAGE TANK
fENT, OVERFLOW
AND DRAIN
-IRA N SFER
PUMP
DILUTION
WATER
SAMPLE TAP
POINT OF
APPL IC All ON
10 - 40,
-------
Recommended Materials
for Use With 50% NaOH
Components Up to 140° F
Rigid Pipe Standard WeigJ it Black Iron
Flexible Connections Rigid Pipe with Ells or Swing Joints,
Stainless Steel or Rubber Hose
Diluting Tees Type 304 Stainless Steel
Fittings Steel
Permanent Joints Wclded or Screwed Fittings
Unions Screwed Steel
Valves — Non-leaking (Plug)
Body Steel
Plug Type 304 Stainless Steel
Pumps (Centrifugal)
Body Steel
Impeller Ni-Resist
Packing Blue Asbestos
Storage Tanks Steel
Pacing and Control. Controls as listed for liquid alum also apply to liquid
caustic soda equipment.
10.5 Carbon Dioxide
Properties and Availability. Carbon dioxide, C0 2 , is available for use m
wastewater treatment plants in gas and liquid form. The molecular weight of CO 2 is
44. Dry CO 2 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 CO 2 in water are very reactive chemically and form
carbonic acid. Saturated solutions of CO 2 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 CO 2 gas only. These generators are usually fired with
combustible gases, fuel oil, or coke and have CO 2 yields as shown in Table 10-4 (8).
Table 10-4
CO 2 YIELDS OF COMMON FUELS
Fuel Quantity CO 2 Yield
(Ib)
Natural Gas 1 ,000 ft 3 115
Coke JIb 3
Kerosene I gal. 20
Fuel Oil (No. 2) 1 gal. 23
Propane 1,000 ft 3 141
Butane 1,000 ft 3 142
10 -41
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The gas forms, as generated at the plant site, usually have a CO 2 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 lb cylinders, 10 to
20 ton trucks and 30 to 50 ton rail cars. The commercial liquid form has a 99.5%
CO 2 content.
Current prices range from $30/ton for 3,000 tons per year and over, to $68/ton for
a quantity of ISO tons/year. These prices include an allowance for freight within a
100 mile radius of the point of manufacture. Another $6/ton may be added for each
additional 100 miles to the point of destination. Major producers of commercial CO 2
are listed in Table 10-5.
Table 10-5
LOCATION OF CARBON DIOXIDE MANUFACTURING PLANTS
Location Manufacturer
CALIFORNIA
Watson (Los Angeles) Liquid Carbonic
Oakland Liquid Carbonic
Brea Airco
Lathrop Airco
Ventura Cardox
Taft Standard Oil
GEORGIA
Augusta Liquid Carbonic
ILLiNOIS
Morris Cardox
Chicago Airco
INDIANA
Jeffersonville Cardox
IOWA
Clinton Airco
Ft. Madison Liquid Carbonic
Ft. Dodge Liquid Carbonic
Muscatine Publicker
KANSAS
Dodge City Liquid Carbonic
Lawrence Airco (I 972)
Lawrence Cardox
KENTUCKY
Doerun (Brandenburg) Olin
10 -42
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Table 1015 (Continued)
Location Manufacturer
LOUISIANA
New Orleans Liquid Carbonic
Luling Airco
MASSACHUSETfS
Tewksbury Liquid Carbonic
MISSISSIPPI
Yazoo City Airco
MISSOURI
Kansas City Airco
Le May (St. Louis) Cardox
NEW JERSEY
Paulsboro Olin
Belleville Liquid Carbonic
Deepwater Airco
NEW MEXICO
Bueyeros SEC
Solana SEC
Mosquero SEC
NEW YORK
Olean Airco
OHIO
Toledo Cardox
Oregon (Toledo) Liquid Carbonic
Lima Airco
Huron Cardox
PENNSYLVANIA
Philadelphia Liquid Carbonic
Thermice (Philadelphia) Publicker
TENNESSEE
Woodstock (Memphis) Cardox
TEXAS
Texas City Liquid Carbonic
Dallas Cardox
Dumas Diamond Shamrock
VIRGINIA
Hopewell Airco
Saltville Olin
10 -43
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Table 10-5 (Continued)
Location Manufacturer
WASHINGTON
Finley Airco
Manufacturers and Addresses
Cardox Div. of Chemetron Corp. Publicker md., Inc.
Dept. TR Walnut & Thomas
840 N. Michigan Avenue Philadelphia, Pa.
Chicago, Illmois 60611 SEC
Airco Industrial Gases Div. of 1033 Humble Place
Air Reduction Co. El Paso, Texas 79987
575 Mountain Avenue
Diamond Shamrock Chemical Co.
Murray Hill, N. J.
300 Union Commerce Building
Liquid Carbonic Corporation Cleveland, Ohio 44115
Dept. TR
135 S. LaSalle Standard Oil Company of California
Chicago, Illinois 60603 225 Bush
San Francisco, California 94104
Olin Corporation
Chemicals Division
745 Fifth Avenue
New York, New York
General Design Considerations. Recovery of CO 2 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 hand 1mg dry air or gas. On-site
generating units have a limited range of CO 2 production as compared with the liquid
storage and feed system, and therefore may require multiple units.
The liquid CO 2 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 be used to feed CO 2 .
Storage. This section applies only to use of commercial liquid CO 2 . Liquid
system capacities encountered in wastewater treatment usually require on-site bulk
storage units. Standard pre-packaged units are available, ranging in size from 3/8 to 50
10 - 44
-------
tons capacity, and are furnished with temperature-pressure controls to maintain
approximately 300 lb/in. 2 at 0°F conditions. The typical package unit contains
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.
Feeding Equipment. Feeding systems for the stack gas source of CO 2 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 CuIp and CuIp (9).
Pressure generators and submerged burners are regulated by valving arrangements on
the fuel and air supply. Generation of CO 2 by combustion is usually difficult to
control, requires frequent operator attention and demands considerable maintenance
over the life of the equipment.
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, the CO 2 with suitable metering and
pressure reduction may be fed directly to the point of application as a gas. However,
vacuum operated, solution type gas feeders are often preferred. Such feeders generally
include safety devices and operating controls in a compact panel housing, with
materials of construction suitable for CO 2 service. Absorption of CO 2 in the injector
water supply approaches 100% when a ratio of 1.0 lb of gas to 60 gal. of water is
maintained.
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
3 I 6 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 decper, to assure complete absorption of the gas.
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 C0 2 , 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 CO 2 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 CO 2 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
CO 2 feed.
10 -45
-------
CO 2 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.
10.6 Polymers
10.6.1 Dry Polymers
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.
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 per cent 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.
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 3 16 stainless steel, FRP, or plastic lined tanks.
Feed 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 10-1 1. It is generally
advisable to keep the holding or storage time of polymer solutions to a minimum, I
to 3 days or less, to prevent deterioration of the product.
Piping and Accessories. Selection must be made after determination of the
polymer, however, type 3 1 6 stainless steel or plastics are generally used.
Pacing and Controls. 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 - 46
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DRY
ISSOLVING—AGING
TAN K
TANK
ION FEEDER
FIGURE 10-lI TYPICAL SCHEMATIC OF
A DRY POLYMER FEED SYSTEM
WATER
I SPERSER
IXER
POINT OF
APPLI CATI ON
10 - 47
-------
HOLDING TANK
—LEVEL PROBE
FIGURE 10-12 TYPICAL AUTOMATIC POLYMER FEED SYSTEItI
HOT
WATER
LEVEL
PROBE
B L EN I
NOTE: CONTROL & INSTRUMENTATION
WIRING IS NOT SHOWN
0
00
TRANSFER PUMP
SOLUTI ON
FEEDERS
POINT OF APPLICATION -
FOR LARGE PLANTS
-------
10.6.2 Liquid Polymers
Properties and Availability. As with dry polymers, there is a wide variety of
products, and manufacturers should be consulted for specific information.
General Design Considerations. Liquid systems differ from the dry systems only
in 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.
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 feeding 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 I mg/I. 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
Plant Size Estimated Cost
(mgd) ($)
5,810
10 26,740
100 203,000
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 tO
nigd 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 Rapid Mixing
For iron or aluminum compounds to be effective in phosphorus removal they must be
intimately mixed with the wastewater Chemical dispersion usually occurs in some type
10 -49
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of mechanically-agitated rapid mix basin. (Mixing may also be accomplished by air
agitation or a hydraulic jump.) Design of these basins is based on detention time and
the velocity gradient in the basin. Walker (12) formulates the velocity gradient, C, as
follows
Where:
w = Water horsepower x 550
Tank Volume (cubic feet)
p = absolute viscosity of the fluid at the
temperature involved, lb-sec/ft 2
Cuip and CuIp (9) recommend a velocity gradient of 300 ft/sec/ft or greater, a
detention time of 15 to 60 seconds, and a power input to the mechanical mixers of
0.25 to 1.0 hp/mgd.
Others indicate that higher values of G, 700 to 1,000 ft/sec/ft are necessary for
efficient blending. To attain a velocity gradient of 800 ft/sec/ft 2 at a detention time
of 30 seconds and 68°F, the water horsepower per mgd can be calculated as follows:
Volume of the mixing basin per mgd = 46.4 ft 3
( HP ) (55 0)/46.4
0.2098 x io
(800)2 x 0.2098 x I o = 550 HP /46.4
HP = 1.13
The required motor (brake) horsepower is:
BHP = HP
Efficiency
Assuming an efficiency of 75%, brake HP = HP /0.75 1.13/0.75 = 1.5. Therefore,
1.5 horsepower per mgd would be required to provide the desired mixing. Using Cuips’
figures of 300 ft/sec/ft, a 30 sec detention time, and the same temperature; the
required HP = 0.16/mgd. Based on the same overall efficiency (75%), the required
brake horsepower would be 0.21/mgd. Full scale studies should be conducted on rapid
mix basins using the higher velocity gradients (700 to I ,000 ft/sec/ft) to investigate
any advantages. The cost of mixers rated at a horsepower capable of producing the
high G values does not appear to be prohibitive.
Recent investigations (13) emphasize the hydrolysis effect on trivalent aluminum and
iron salts by dilution. Solutions of these salts will not hydrolyze as long as the solution
10 - 50
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pH is 3.0 or less. Upon dilution to a higher pH, hydrolysis occurs; that is, the trivalent
aluminum and iron form hydrous oxide gels. Following this, the ions age through a
series of trivalent species by polymerization until the neutral hydroxide form is
attained. Hydrolysis takes place in less than 1/10 second and the neutral trivalent
hydroxide species is reached within two or three minutes. Phosphorus precipitation
capability of the aluminum or iron trivalent ion decreases as hydrolysis and aging
progress; therefore, it is imperative to instantaneously disperse the chemical in the
wastewater. The coagulant should be added to the vortex under a mechanical mixer or
to the point of greatest turbulence in other types of mixers. This will assist in
achieving the most rapid dispersion possible.
10-51
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10.8 References — Chapter 10
1. ________________ , “Chemicals Used in Water and Wastewater Treatment-Aluminum
Sulfate”, Wat and Wastes Eng., Dec. 6:12 p46(1969).
2. ________________ , “Chemicals Used in Water and Wastewater Treatment-Ferric
Chloride”, Wat and Wastes Eng., 7:3, p 65 (1970).
3. _________________ , “Chemicals Used in Water and Wastewater Treatment-Ferric
Sulfate”, Wat and Wastes Eng, 7:4, p 69(1970).
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, Rhode Island (June, 1969).
6. National Lime Association, Personal Communication (May, 1971).
7. “Caustic Soda”, PPG Industries, Inc., Chemical Division, Pittsburgh, Pa. (1969).
8. Haney, P. D. and Hamann, C. L., “Recarbonation and Liquid Carbon Dioxide”,
JAWW 1 4, 61:10, p 512 (1969).
9. CuIp, R. L., and Cuip, G. L., Advanced Wastewater Treatment, Van Nostrand-
Reinhold Company, New York (1971).
10. Carr, R. L., “Polyelectrolyte Coagulant Aids-Dry and Liquid Handling and
Application”, Wat. and Sew. Works — Reference No., I l4:R.N., p 4-64, (1967).
II. Russo, F. and Carr, R. L., “Polyelectrolyte Coagulant aids and Flocculants: Dry and
Liquid, Handling and Application”, Wat. and Sew Works — Reference No,
l17:R.N., p R-72 (1970).
12. Walker, J. D., “High Energy Flocculation Units”, JAWWA, 60:11, p 1271 (1968).
13. Walker, J. D., Personal Communication (July, 1971)
10.9 Supplemental Bibliography
10.9.1 General
Water Quality and Treatment, Third Edition, American Water Works Association, inc.,
McGraw-Hill, New York (1971).
Water Treatment Plant Design, American Society of Civil Engineers, American Water
Works Association, and Conference of State Sanitary Engineers, New York (1969).
Ockershausen, R. W., “Safe Handling of Water Works Chemicals”, JA WWA, 63:6, p 336
(1971).
10 - 52
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“Chemicals Used in Treatment of Water and Wastewater”, BIF Reference No. 1.21-15,
BIF industries, Providence, Rhode Island (1970).
10.9.2 Aluminum Compounds
Truitt, V., “We Converted to Liquid Alum Treatment”, Public Works, 99:11, p 88
(1968).
“Aluminum Sulfate”, Allied Chemical Corporation, Morristown, New Jersey (1968).
“Alum”, American Cyanamid Company, Wayne, New Jersey (1964).
“Aluminum Sulfate”, Stauffer Chemical Company, New York (1967).
10.9.3 Iron Compounds
“Handling Ferric Chloride”, Dow Chemical Company, Midland, Michigan (1965).
“Pennsalt Ferric Chloride”, Pennsalt Chemicals Corporation, Philadelphia, Pennsylvania
(1965).
10.9.4 Lime
“Chemical Lime Facts”, National Lime Association, Washington, D. C. (1964).
Carr, R. L., “Limestone to Lime to Slaked Lime”, Wat. and Sew. Works — Reference
No, 1 13:R.N., p R-61 (1966).
“Chemicals Used in Water and Wastewater Treatment — Calcium
Hydroxide (Hydrated Lime)”, Wat and Wastes Eng., 7:1, p 53 (1970).
“Chemicals Used in Water and Wastewater Treatment — Quicklime”,
Wat. and Wastes Eng., 7:5, p 52 (1970).
10.9.5 Soda Ash
“Soda Ash”, Allied Chemical Corporation, New York (1966).
“What’s Different about FMC Soda Ash?”, FMC Corporation, New York (1964).
“Stauffer ‘Natural’ Soda Ash”, Stauffer Chemical Company, New York (1964).
“Wyandotte Soda Ash”, Wyandotte Chemicals Corporation, Wyandotte, Michigan (1955).
10.9.6 Caustic Soda
“Caustic Soda”, PPG Industries, Inc., Chemical Division, Pittsburgh, Pennsylvania (1969).
“Unloading Liquid Caustic from Tank Cars”, Manufacturing Chemists Association,
Washington, D. C. (1968).
10 - 53
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“Properties and Essential Information for Safe Handling and Use of Caustic Soda”,
Manufacturing Chemists Association, Washington, D.C. (1968).
“Caustic Soda”, Allied Chemical Corporation, New York (1963).
“Caustic Soda Handbook”, Diamond Alkali Company, Cleveland, Ohio (1967).
“Dow Caustic Soda Handbook”, Dow Chemical Company, Midland, Michigan (Undated).
“Hooker Caustic Soda”, Hooker Chemical Corporation, Niagara Falls, New York (1966).
“Olin Caustic Soda”, Olin Mathieson Chemical Corporation, New York (1961).
“Pennsalt Caustic Soda”, Pennsalt Chemicals Corporation, Philadelphia, Pennsylvania
(1964).
“Caustic Soda”, Stauffer Chemical Company, New York (1965).
“Caustic Soda by Wyandotte”, Wyandotte Chemicals Corporation, Wyandotte, Michigan
(1961).
10 - 54
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Chapter 11
CHECKLiSTS FOR REVIEW OF PLANS AND SPECIFICATIONS
11.1 introduction
The processes outlined in this design manual have been developed for application to
projects where improved effluent quality is needed or required. In some cases, these
processes have had limited use in full-scale design of wastewater treatment facilities
Design engineers and reviewing authorities may not be completely familiar with the
parameters to be considered in the preparation of plans and specifications for a given
project. The following plan and specification review check sheet has been prepared to
serve as a guide to the engineer in the design of proposed facilities which use the
process or processes outlined in this manual. It will also be used by the Environmental
Protection Agency, and may be used by State and local authorities in their review of
projects for approval.
As with the case of any check sheet, its purpose is to fully consider all possible
parameters for an individual process. For a given project all or part of the check sheet
may be applicable and should be used with this fact in mind.
11 .‘2 Phosphorus Removal by Mineral Addition Before the Primary Settler
11.2.1 Rapid Mix
Influent flow, mgd
Number of units
Detention time, seconds
Velocity gradient — G, ft/sec/ft
Mixer horsepower, hp
Freeboard, ft
11.2.2 Flocculator
Flow and recirculation, mgd
Number of units
Detention time, minutes
Drive horsepower, hp
Velocity gradient — G, ft/sec/ft
Velocity gradient x time — Gt
Freeboard, ft
11.2.3 Primary Settler
Flow and recycle, mgd
Number of units
Detention time, hours
Surface overflow rate, gpd/ft 2
11—1
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Weir rate, gpd/linear ft
Sidewater depth, ft
Freeboard, ft
Solids removal efficiency, %
Clarifier mechanism loading, lb/linear ft
Solids pumping capacity, gpm
11.2.4 Chemical Storage and Feed System
Dosage rate, mineral : phosphorus (molar ratio)
Dosage rate, mg/I
Bulk storage
Capacity
Construction material
Pumps, transfer
Type
Capacity, gpm
Rated head, ft
Dilution and feed tanks
Capacity, gal.
Construction material
Agitators
Agitator horsepower, hp
Construction material
Feeder or feed pumps
Type
Capacity, gph or lb/hour
Number of operating units
Number of standby units
Control (flow, influent P)
11.2.5 Sludge Handling
Sludge lines — capacity and head loss, gpm and ft
Sludge pumps — capacity and head, gpm and ft
11.2.6 Gravity Thickener
Number of units
Solids loading, lb/ft 2 /day
Overflow rate, gpd/ft 2
Detention time, hours
Sidewater depth, ft
Freeboard, ft
Feed concentration, % dry solids
Thickened concentration, % dry solids
Solids recovery, %
I I - 2
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Thickening mechanism — type
Mechanism loading, lb/linear ft
1 1.2.7 Centrifuge
Number of units
Type
Feed concentration, % dry solids
Underfiow sludge concentration, % dry solids
Solids recovery, %
Dry solids capacity, lb/day
Total sludge, lb/day
Centrate return
11.2.8 Vacuum Filter
Number of units
Type
Feed concentration, % dry solids
Filter cake, % dry solids
Solids recovery, %
Dry solids capacity, lb/hour/ft 2
Total dry solids, lb/day
Filtrate return
11.2.9 Incinerator
Number of units
Sludge moisture, %
Sludge water, lb/hour
Sludge dry solids, lb/hour
Sludge total, lb/hour
Sludge volatile matter, % dry solids
Sludge ash content, % dry solids
Sludge heating value, Btu/lb
Allowable combustibles in ash, %
Operating schedule, hours/week
Auxiliary fuel
Maxim urn incinerator temperature, °F
Air pollution control devices
Maximum stack g s temperature, °F
11.3 Phosphorus Removal by Lime Addition Before the Primary Settler
11.3.1 Rapid Mix
Sec Scctioii 1121
II . 3
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11.3.2 Flocculátor
See Section I I 2.2
11.3.3 Primary Settler
See Section 11.23
1 1.3.4 Lime Storage and Feed System
Dosage rate, mg/I
Bulk storage
Capacity, tons
Construction material
Bin dimensions, ft
Dust collector
Lime feeders
Number of feeders
Type
Maximum capacity, lb/hr
Accessories
Lime slaker
Number oI siakers
Type
Maximum capacity, lb/hr
Slaking time, minutes
Lime slurry holding tank
Capacity, gal
Agitator horsepower, hp
Lime slurry pumps
Type
Capacity, gpm
Number of operating units
Number of standby units
Control (pH, flow)
11.3.5 Sludge Handling
See Section II 2 5
11.3.6 Gravity Thickener
See Section I I 2.6
I L3.7 Centrifuge
See Section 11.2.7
11.3.8 Vacuum Filter
See Section 11.2.8
11 -4
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11.4 Phosphorus Removal in Trickling Filters by Mineral Addition
11.4.1 Rapid Mix (if necessary)
See Section 11.2.1
11.4.2 Flocculator (if necessary)
See Section 11.2.2
11.4.3 Final Settler
See Section 11 2.3
11.4.4 Chemical Storage and Feed System
Point of Chemical Addition
(See Section 11 2.4 for remaining parameters)
11.4.5 Sludge Handling
See Section 11 2 5
11.4.6 Gravity Thickener
See Section 11 2.6
11.4.7 Centrifuge
See Section 11.2.7
11.4.8 Vacuum Filter
See Section 11.2.8
11.4.9 Incinerator
See Section 11.2.9
11.5 Phosphorus Removal in Activated Sludge Plants by Mineral Addition
11.5.1 Rapid Mix (if necessary)
See Section I I .2 I
11.5.2 Flocculator (if necessary)
See Section 11 2.2
11.5.3 Final Settler
See Section 1 I 2.3
11 -5
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11.5.4 Chemical Storage and Feed System
Point of Chemical Addition
(See Section 11.2.4 for remaining parameters)
11.5.5 Sludge Handling
See Section 11 .2.5
11.5.6 Gravity Thickener
See Section I 1 2.6
11.5.7 Centrifuge
See Section 11.2.7
11.5.8 Vacuum Filter
See Section 11.2 8
11.5.9 Incinerator
See Section I I .2.9
11.6 Phosphorus Removal by Lime Treatment of Secondary Effluent
11.6.1 Rapid Mix
See Section 11.2 1
11.6.2 Flocculator
See Section I I .2 2
11.6.3 Tertiary Settler
See Section I 1 .2 3
11.6.4 Recarbonator
Number of stages
Basins
Detention time, minutes
Sidewater depth, ft
Inside depth, ft
CO 2 source
CO 2 dosage, Ib/lO 6 gat.
CO 2 dispersion
Location in basins
Type
Control
11 -6
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Depth of submergence, ft
Size of diffuser orifices, in.
1 1.6.5 Multimedia Filter
Average flow at steady state, gpm
Filtration rate, gpm/ft 2
Total surface area, ft 2
Available head, ft
Media
Type
Depth, in.
Effective size, mm
Uniformity coefficient
Support material
Underdrains
Type
Head loss, ft
Freeboard, ft
Backwash
Maximum rate, gpm/ft 2
% Expansion
Length of cycle, minutes
Washwater troughs
Number
Distance from top of media to bottom of trough, ft
Surface wash
Rate, gpm/ft 2
Duration, minutes
Air wash
Rate ft 3 /minute/ft 2
Air pressure, lb/in 2
Duration, minutes
I 1.6.6 Lime Storage and Feed System
See Section 11.3.4
11.6.7 Sludge Handling
See Section I 1 2 5
11.6.8 Gravity Thickener
See Section 11.2 6
11.6.9 Centrifuge
See Section 11 .2 7
11 -7
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11.6.10 Vacuum Filter
See Section II 2.8
11.6. 11 Recalcining Furnace
Number of units
Type
Operating schedule, hours/week
Furnace diameter, ft
Number of hearths
Composition of the feed sludge
Volatile suspended solids, %
Ash, %
Fuel value of volatile suspended sol]ds, Btu/lb
Sludge concentration, % dry solids
Dry solids, lb/day
Water, lb/day
Auxiliary fuel
Maximum recalciriation temperature, °
Air pollution control devices
Maximum stack gas temperature, °F
Recalcined lime cooler
11.6.1 2 Recalcined Lime Storage
Capacity, tons
Construction material
Bin dimensions, ft
Dust collector
11.7 Phosphorus Removal by Mineral Addition to Secondary Effluent
11.7.1 Rapid Mix
See Section I I 2 1
11.7.2 Flocculator
See Section 11.2.2
11.7.3 Tertiary Settler
See Section 11.2.3
11.7.4 Multimedia Filter
See Section Il 6.5
1 1 - 8
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11.7.5 Chemical Storage and Feed System
See Section 11.2.4
11.7.6 Sludge Handling
See Section 11.2.5
11.7.7 Gravity Thickener
See Section 11 2 6
11.7.8 Centrifuge
See Section 11.2.7
11.7.9 Vacuum Filter
See Section 11 2 8
11.7.10 Incinerator
See Section 11.2.9
11 -9
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APPENDIX
Abbreviations
The following are abbreviations of terms and dimensions used in the manual
Terms
BOD Biochemical oxygen demand
BOD 5 Five day BOD
COD Chemical oxygen demand
DO Dissolved oxygen
FRP Fiberglas reinforced plastic
MLSS Mixed liquor suspended solids
SS Suspended solids
VSS Volatile suspended solids
Dimensions
gpd Gallons per day
gph Gallons per hour
gpm Gallons per minute
hp Horsepower
JTU Jackson turbidity units
mgd Million gallons per day
rpm Revolutions per minute
A-I
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ACKNOWLEDGMENT
This Manual was prepared by Black & Veatch, Consulting Engineers under the sponsor-
ship of the U. S. Environmental Protection Agency Contributions were made by
Shimek-Roming-Jacobs & Finklea Consulting Engineers, The Dow Chemical Company,
and the Environmental Protection Agency staff Their assistance during the preparation
of the manual is gratefully acknowledged.
AC-I
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_ 1 Titio
PROCESS DESIGN MANUAL FOR PHOSPHORUS REMOVAL
101 Auth (Ø
Black & Veatch, Consulting
Engineers
PiuJoaS D.si*lati
17010 GNP
!JAvajIable from the U. S. Environmental Protection
Agency, Office of Water Programs, Technology
Transfer, Washington, D. C.
2 21 cht 1
23j Desc 1ptore ‘SCaZred Fire()
*Phosphorus, *Chemical precipitation, *eutrophication, *coagulation, *f].occulation,
*gettling basins, *filtration, *mixing, dewatering, sludge treatment, solids contact
process, tertiary treatment, municipal wastes, wastewater treatment.
251 Identifier. (Starred First)
*Phosphorus removal, *alum, *ferrous, *ferric, *lime treatment, sodium aluminate,
pickle liquor, ,dual media, multimedia, thickening, advanced waste treatment.
271 Abstract
The discharge of phosphorus—containing vastewaters 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 precip-
itation 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 parani ters, and
costs. This information should be of value to designers, municipal officials,
regulatory agencies, city planners, and treatment plant operators.
Abstracter Dr. Carl A. Brunner
WR 105 (REV. JULY l9 9
WRSI C
[ IflZtitIitIotl Robert A. Taft Water Research Center
SEND. WITH COPY OF DOCUMENT. TO: WATER RESOURCES SCIENTIFIC INFORMATION CENTER
U.S. DEPAATMENT OF THE INTERIOR
WASHINGTON. D. C. 2CS40
U. S. Environmental Protection Agency
SELECTED WATER RESOURCES ABSTRACTS
INPUT TRANSACTION FORM
* U 5 GOVESNUINT WIlTING 0Ui l 1571 0— 4 4 1 — 507
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