United States
Environmental Protection
Agency
Municipal Environmental Research
Laboratory
Cincinnati OH 45268
EPA-600/2-73-V2
Research and Development
4>EPA
Operational Results
for the Piscataway
Model 5 MGD
AWT Plant
-------
RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S. Environmental
Protection Agency, have been grouped into nine series. These nine broad cate-
gories were established to facilitate further development and application of en-
vironmental technology. Elimination of traditional grouping was consciously
planned to foster technology transfer and a maximum interface in related fields.
The nine series are:
1 Environmental Health Effects Research
2. Environmental Protection Technology
3. Ecological Research
4. Environmental Monitoring
5. Socioeconomic Environmental Studies
6. Scientific and Technical Assessment Reports (STAR)
7. Interagency Energy-Environment Research and Development
8. "Special" Reports
9. Miscellaneous Reports
This report has been assigned to the ENVIRONMENTAL PROTECTION TECH-
NOLOGY series. This series describes research performed to develop and dem-
onstrate instrumentation, equipment, and methodology to repair or prevent en-
vironmental degradation from point and non-point sources of pollution. This work
provides the new or improved technology required for the control and treatment
of pollution sources to meet environmental quality standards.
This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.
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EPA-600/2-78-172
September 1978
OPERATIONAL RESULTS FOR THE
PISCATAWAY MODEL 5 MGD AWT PLANT
by
Thomas P. O'Farrell
Construction Operations Branch
Office of Water Programs
Washington, D.C. 20460
Robert A. Menke
Washington Suburban Sanitary Commission
Hyattsville, Maryland 20781
Grant No. S-802943 .
Project Officers
Thomas P. O'Farrell
Washington, D.C. 20460
D. F. Bishop, F. L. Evans, S. A. Hannah
Wastewater Research Division
Municipal Environmental Research Laboratory
Cincinnati, Ohio 45268
MUNICIPAL ENVIRONMENTAL RESEARCH LABORATORY
OFFICE OF RESEARCH AND DEVELOPMENT
U.S. ENVIRONMENTAL PROTECTION AGENCY
CINCINNATI, OHIO 45268
-------
DISCLAIMER
This report has been reviewed by the Municipal Environmental Research
Laboratory, U.S. Environmental Protection Agency, and approved for publi-
cation. Approval does not signify that the contents necessarily reflect
the views and policies of the U.S. Environmental Protection Agency, nor
does mention of trade names or commercial products constitute endorsement
or recommendation for use.
11
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FOREWORD
The Environmental Protection Agency was created because of increasing
public and government concern about the dangers of pollution to the health
and welfare of the American people. Noxious air, foul water, and spoiled
land are tragic testimony to the deterioration of our natural environment.
The complexity of that environment and the interplay between its components
require a concentrated and integrated attack on the problem.
Research and development is that necessary first step in problem solu-
tion and it involves defining the problem, measuring its impact, and search-
ing for solutions. The Municipal Environmental Research Laboratory develops
new and improved technology and systems for the prevention, treatment, and
management of wastewater and solid and hazardous waste pollutant discharges
from municipal and community sources, for the preservation and treatment of
public drinking water supplies, and to minimize the adverse economic, social,
health, and aesthetic effects of pollution. This publication is one of the
products of that research; a most vital communications link between the
researcher and the user community.
Conventional treatment of municipal wastewater produces an effluent
that may need additional treatment if a high quality effluent is required
for discharge or reuse. A number of tertiary treatment processes have
been developed and evaluated at laboratory and small pilot scale but at
the inception of this project, had not been adequately demonstrated at full
scale. This publication reports the performance of a 5 mgd tertiary treat-
ment plant using lime clarification, dual media filtration and granular
activated carbon adsorption.
Francis T. Mayo, Director
Municipal Environmental Research
Laboratory
111
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ABSTRACT
A 5 mgd tertiary wastewater treatment plant was constructed to
demonstrate treatment of effluent from a 5 mgd step aeration activated
sludge plant. The two-stage high lime process with intermediate re-
carbonation, filtration and activated carbon adsorption operated at
the design rate for 36 days between two failures of the reactor clari-
fiers. A single-stage low lime process with filtration and activated
carbon adsorption operated for 89 days. The combined secondary and
tertiary treatment removed > 97% of BOD, TSS and P in the raw waste-
water. Capital cost of the 5 mgd two-stage high lime system was 4.7
million dollars and operating costs were estimated as 36 cents per
1000 gallons of wastewater.
IV
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CONTENTS
Foreword iii
Abstract iv
Figures vi
Tables viii
Conversion Factors xii
Acknowledgements xiii
I. Introduction 1
II. Conclusions 3
III. Detailed Description of Secondary Treatment
Facility 4
IV. Tertiary Treatment 8
V. Detailed Description of the Model Tertiary
Plant Facility 15
VI. Results of Two-Stage High Lime Evaluation 37
VII. Results of Single-stage Low Lime Evaluation .... 53
VIII. Carbon Regeneration 65
IX. Cost Analysis 69
Appendix • 88
v
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FIGURES
Number Paqe
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
Flow Schematic of the Piscataway Secondary Plant
Schematic of the 5 mgd System of the Piscataway
Secondary Plant
Two-Stage High Lime Tertiary Process
Single-stage Low Lime Tertiary Process
Lime Handling System
Cross Section of Reactor Clarifier
Solid Bowl Centrifuge Section
Centrifuge Operation for Total Capture
Centrifuge Operation for Wet Classification
Cross Section of Multiple Hearth Furnace
Cross Section of Dual Media Filter
Flow Schematic for Carbon Regeneration
Cross Section of Carbon Adsorber Underdrain
Comparison of Alkalinity S TKN of Secondary Effluent...
Distribution of Operating Costs for the Low Lime Process
with Wasting of Wet Solids
Distribution of Operating Costs for the Low Lime Process
with Solids Dried and Wasted
Distribution of Operating Costs for the High Lime Process
Distribution of Capital Costs for the Low Lime Process..
5
6
10
11
21
23
26
27
28
30
32
35
36
39
78
79
80
81
VI
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FIGURES (CONTINUED)
Number Paqe
19 Distribution of Capital Costs for the High Lime
Process 82
20 Distribution of Power Requirements for the Low
Lime Process 83
21 Distribution of Power Requirements for the High
Lime Process 84
VI1
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TABLES
Number Page
1 Design Data for Model Plant Equipment 16
2 Model Plant Pumps 19
3 Major Equipment Vendors 20
4 Operating Conditions of the Piscataway Secondary Plant
During the High Lime Process Evaluation 38
5 Removal of Biochemical Oxygen Demand (BOD 5 Day) During
the High Lime Process Evaluation 38
6 Removal of Suspended Solids During Evaluation of the
High Lime Process 40
7 Removal of Total Phosphorus (as P) During Evaluation
of the High Lime Process 40
8 Removal of Nitrogen Compounds During Evaluation of the
High Lime Process 41
9 Loading Rates During Evaluation of the High Lime Process. 41
10 Plant Recycle During High Lime Evaluation 42
11 Chemical Usage in the High Lime Process 44
12 Removals of Chemical Oxygen Demand (COD) and Total
Organic Carbon (TOC) During Evaluation of the High
Lime Process 45
13 Performance of Carbon Adsorber Train #1 During
Evaluation of the High Lime Process 47
14 Performance of Carbon Adsorber Train #2 During
Evaluation of the High Lime Process 48
15 Performance of Carbon Adsorber Train #3 During
Evaluation of the High Lime Process 49
Vlll
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TABLES (CONTINUED)
Number Page
16 Cumulative COD and TOG Loadings on Carbon at the end
of the 36-Day High Lime Evaluation 50
17 Solids Material Balances for the High Lime Evaluation... 50
18 Operating Conditions of the Piscataway Secondary
Operation During the Low Lime Process Evaluation 54
19 Removal of Biochemical Oxygen Demand (BOD 5 Day) During
Evaluation of the Low Lime Process 54
20 Removal of Suspended Solids During Evaluation of the
Low Lime Process 55
21 Removal of Total Phosphorus (as P) During Evaluation
of the Low Lime Process 55
22 Removal of Nitrogen Compounds During Evaluation of
the Low Lime Process 56
23 Loading Rates During Evaluation of the Low Lime Process.. 57
24 Plant Recycle Flows During Low Lime Evaluation 58
25 Chemical Usage in the Low Lime Process 58
26 Removals of Chemical Oxygen Demand (COD) and Total
Organic Carbon (TOC)| During Evaluation of the Low
Lime Process 59
27 Performance of Carbon Adsorber Train #1 During Evaluation
of the Low Lime Process 61
28 Performance of Carbon Adsorber Train #2 During Evaluation
of the Low Lime Process 62
29 Performance of Carbon Adsorber Train #3 During Evaluation
of the Low Lime Process 63
IX
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TABLES (CONTINUED)
Number Paqe
30
31
32
33
34
35
36
37
38
39
40
41
42
43
44
45
46
47
Cumulative COD and TOG Loadings at the End of the
Daily Solids Production for the Low Lime Evaluation . .
Inventory of Carbon .................................
Operating Data for Regeneration of Carbon. ...........
Results of Laboratory Analyses of Carbon. ............
Carbon Loadings at Time of Regeneration
Sieve Analyses of Carbon from T-18 Carbon Adsorber...
Furnace Conditions During Carbon Regeneration. .......
Capital Costs of the Model Plant
Distribution of Capital Costs of the Model Plant
WSSC Construction Phases for the Piscataway
Secondary Treatment Plant
Cost Breakdown for the 30 mgd Secondary Treatment Plant
Costs of Engineering Services for the Model Plant
Capital Costs for the Model Plant Unit Processes
Breakdown on Capital Costs for the Model Plant
Capital Cost Breakdown for the Model Plant Equipment —
Cost Fiqures for Energy and Chemicals
64
64
65
66
67
67
68
68
70
71
7?
73
74
74
75
76
85
85
X
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TABLES (CONTINUED)
Number Page
48 Operating Costs for the Low Lime Process with
Wasting of Wet Solids 86
49 Operating Costs for the Low Lime Process with
Solids Dried and Wasted 86
50 Operating Costs for the High Lime Process 87
51 Personnel Breakdown by Unit Processes 87
XI
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CONVERSION FACTORS
The following factors convert the English units used in this report
to the SI metric unit in popular usage in water engineering practice.
English Unit X
Acre
Btu
cu ft
OF
ft
gal
gpm
gpm/sq ft
hp
in
Ib
mil gal
ton
Multiplier
4,047
1.055
0.028
0.555(°F-32)
0.3048
3.785
0.0631
40.7
0.7454
2.54
0.454
3,785
907
= Metric Unit (SI)
m2
kJ
m3
°C
m
£
Vsec
£/min m2
kW
cm
kg
m3
kg
Xll
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ACKNOWLEDGEMENTS
We express our sincere appreciation to Herman Keys, Senior Plant
Operator, and his staff of Commission Plant Operators for their dedication
and initiative during start up and testing periods.
We thank the Commission's Herb Johnson, Jr., and other Electrical
Maintenance Personnel for their start up assistance and maintenance
service.
Also appreciation to Raul Celerio, Chief Chemist, and his laboratory
personnel for testing and analysis of process samples and special gas tests,
We thank Lam Lim and Dave Thorne, Supervisors of the Secondary Plant
for their continuous involvement with the plant and the flow coming to the
Model Plant.
EPA's Blue Plains Pilot Plant should also be mentioned for sample
analysis and technique support during the testing period.
xiii
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I. INTRODUCTION
The WSSC Wastewater Treatment Facility is located in Piscataway, Mary-
land, south of Washington D. C., at the inlet to Piscataway Bay. The efflu-
ent from the plant discharges into the marshlands of Piscataway Bay and as
a result, high quality effluent is required. Ultimately the effluent from
the facility will flow via an underground pipeline across federally owned
property and discharge into the deep water of the Potomac River. To date,
construction of the pipeline has not been completed and the WSSC has made a
maximum effort to reduce the load to Piscataway Bay. The secondary treat-
ment facilities at Piscataway have undergone three major construction phases
which include:
1. Construction of a 5 mgd step aeration activated sludge system in 1967.
2. Completion of 28.6 acres of polishing ponds in 1970.
3. Expansion of the step aeration system to 30 mgd and the installation
of fluid solids incinerators in 1974.
The plant was first placed in service in November 1967.
Early in 1966, the Environmental Protection Agency, formerly the Federal
Water Pollution Control Administration, and the Washington Suburban Sanitary
Commission entered into a joint agreement for the construction and operation
of a tertiary treatment facility. The "Joint FWPCA-WSSC Model Advanced Waste
Treatment Plant" was to be located in Piscataway, Maryland, at the 313 acre
site of the WSSC secondary facilities. The purpose of the plant was to demon-
strate by advanced waste treatment techniques the removal of COD, BOD, sus-
pended solids, phosphorus and refractory organics from secondary municipal
wastewaters.
A 5 mgd step aeration activated sludge plant was under construction and
was scheduled to be completed in late 1967 at the Piscataway site. The plant
was to treat a mainly domestic wastewater from the rapidly expanding Prince
Georges County, Maryland. The wastewater is pumped from sanitary sewers
located in the Broad Creek, Swan Creek and Piscataway Creek drainage areas
in Southern Prince Georges County.
In addition to the available plant site, the decision to construct the
plant at Piscataway was also based on the increased interest in reducing the
pollution load to the Potomac River Estuary. The Research and Development
Program of the FWPCA had demonstrated in the laboratory and at small pilot-
plant scale the technical feasibility of improving carbon and phosphorus re-
movals from municipal waste discharges. In 1967, the only tertiary treatment
system in operation was the 2.5 mgd facility at Lake Tahoe, California.
-------
At that plant, alum was being used for chemical clarification. It was de-
cided that full-scale operation was necessary to determine process relia-
bility and to secure accurate cost information for the construction and
operation of a tertiary treatment plant.
In December, 1966, a grant (WPRD 62-01-67) was awarded to the Washington
Suburban Sanitary Commission by the FWPCA for the design and construction of
a 5 mgd full-scale tertiary treatment plant. The plant was conceived to
include: chemical clarification with lime, recarbonation, dual-media filtra-
tion, activated carbon adsorption, lime recovery, and activated carbon re-
generation. The total estimated project cost was 2.2 million dollars with
a Federal share of 1.65 million dollars.
Affirmative action on the project was delayed until July 1968, when a
design entitled "Process Design for the Model Advanced Waste Treatment Plant-
Piscataway, Maryland," was forwarded to WSSC by the FWPCA. The design in-
cluded the two-stage high lime process with intermediate recarbonation and
lime recovery by thermal recalcination, filtration, and activated carbon
adsorption with thermal regeneration. Updated information, which had been
obtained at the EPA-DC Pilot Plant in Washington, D.C., treating similar
low alkalinity wastewater, was included in the design. At that time it was
anticipated that the previously estimated 2.2 million dollars would be in-
sufficient to cover the costs of a 5 mgd Advanced Waste Treatment Plant.
Shortly after the process design was submitted to the WSSC, a design engi-
neer w.as selected by WSSC with instructions to prepare a preliminary con-
struction design, based on the submitted R&D process design, and a cost
estimate. In January 1969, the engineer estimated the cost of the project
at 3.2 million dollars. In June 1969, a supplementary grant (17080 DZY)
was awarded to WSSC for $750,000, thus increasing the Federal share to 2.4
million dollars. Because of the limited available funds, the engineer was
instructed to reduce costs where possible. The area most affected by the
cost reduction was the elimination of duplicate equipment that would be
necessary to ensure continuous operation.
The size of the plant was selected as 5 mgd. The final construction
drawings and specifications were completed in May 1970. Review of the bid
responses in September 1970, showed that the lowest bid was approximately
4.5 million dollars. An agreement was reached with the EPA Region III
Construction Grants Division whereby Federal costs in excess of 2.4 million
dollars would be paid by the Construction Grants Division as part of Project
WPC-Md-233. Contracts for the construction of the 5 mgd tertiary facility
were awarded in November 1970, with an expected completion date of January
1972. The final construction and initial operation of the plant were, how-
ever, delayed by many factors.
In April 1972, an Environmental Protection Agency grant was awarded to
the WSSC for one year of operation of the tertiary facility. It was antici-
pated that the system would operate from July 1972, through June 1973.
However, because of construction delays, start-up of the operation was delay-
ed until January 1973. Mechanical failures in the system initially prevented
continuous operation, thus allowing the operating grant to be extended until
June 1974, without additional funds.
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II. CONCLUSIONS
This publication reports results from the operation of a 5 mgd tertiary
wastewater treatment plant used to upgrade the quality of effluent from a
conventional secondary plant using step aeration activated sludge. The
tertiary plant used either two-stage high lime or single-stage low lime
followed^ by dual-media filtration and granular activated carbon adsorption.
Major conclusions from the demonstration project are as follows:
1. The high lime process using an average dosage of 257 mg/1 CaO and
18 mg/1 Fed3 significantly reduced residuals of BOD, TSS and P in the
secondary effluent. BOD was reduced from 16.5 mg/1 to 4.0 mg/1; TSS was
reduced from 27.5 mg/1 to 2.5 mg/1 and P was reduced from 3.50 mg/1 to
0.10 mg/1.
2. The low lime process using an average dosage of 113 mg/1 CaO and
25 mg/1 FeCl3 produced removals of BOD, TSS and P comparable to those obtain-
ed with the high lime process.
3. Tertiary treatment did not significantly affect total N residuals
in the secondary effluent.
4. Carbon losses from regeneration of three columns under less than
optimum conditions were estimated as 8-10%.
5. Operating costs for tertiary treatment in this demonstration plant
were in the range of 29-36 cents per 1000 gallons. These are considered to
be unusually high because data are based on the startup period when the plant
was not at optimum efficiency and operators were not familiar with the plant.
6. A highly competent staff is required to successfully operate a
complex tertiary wastewater treatment plant. Extra efforts should be made
to select, train and retain personnel.
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III. DETAILED DESCRIPTION OF SECONDARY TREATMENT FACILITY
The primary and secondary treatment facilities at Piscataway consist of
two parallel systems (capacity = 5 and 25 mgd) with a common solids handling
system. Raw wastewater is presently pumped to the plant from four pumping
stations with the total capacity of 75 mgd. Although the total feed system
has the capacity to feed 75 mgd, neither the sewage taps nor the plant has
the capacity to collect or treat this volume of wastewater.
The flow enters the plant at the distribution structure via two force
mains. The recycle of overflows and filtrate from the solids handling system
also enters the distribution structure. From the distribution structure,
the flow is split for delivery to three aerated grit chambers. A schematic
diagram of the liquid treatment facilities is shown in Figure 1. Although
the entire secondary plant is presented, only those unit processes associ-
ated with the 5 mgd system feeding the Model Plant will be discussed. The
5 mgd secondary system as operated during the grant is shown in Figure 2.
The raw sewage is manually split and fed to the three grit chambers.
The effluent from the single grit chamber for the 5 mgd system passes through
two 3/8" barminutors equipped with automatic rotating cutters. Four cen-
trifugal pumps, each rated at 1750 gpm, are piped to the open channel follow-
ing the barminutors to provide a constant flow of 5 mgd through the 5 mgd
secondary section into the Model Plant. Since the three chambers are'inter-
connected, all flow in excess of 5 mgd is diverted to the two grit chambers
in the 25 mgd system.
The degritted effluent is split at the primary inlet well and fed to
two parallel primary settling tanks. Each tank with an outside diameter of
60 ft is equipped with both bottom scrapers and surface skimmers. At a flow
of 2.5 mgd per tank, the loading to each tank is 890 gal/day/sq ft with a
detention time of 1.7 hours, The primary solids are wasted by gravity to
the primary solids collection well. At the collection well, the primary
solids are combined with the scum from the surface skimmers and flow by
gravity to the sludge thickener. Wasting of primary solids is a manual
operation without automatic flow measurement.
Primary effluents from the two clarifiers are combined and flow by
gravity to the step aeration basins. Two parallel reactors are provided with
a common feed channel located in the center of the two reactors as shown in
Figure 2. Primary effluent is fed to the quarter points of the reactors at
a rate of Q/6 through each gate. Settled solids from the secondary clari-
fiers are recycled to the head of the two reactors to maintain an average
MLSS concentration of 2000 mg/1." Normally, the recycle flow was maintained
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RAW WASTE WATER-
PLANT RECYCLE —
25 mgd SYSTEM
DISTRIBUTION STRUCTURE
5 mgd SYSTEM
T^> . n
GRIT CHAMBERS (3)
T
LIFT PUMPS (4)
PRIMARY CLARIFIERS (4)
AERATION BASINS (4)
o/
SECONDARY CLARIFIERS (4)L
por
IDS
MODEL PLANT
CHLORINATION BASIN
PISCATAWAY BAY
Figure 1. Flow schematic of the Piscataway secondary plant.
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RAW WASTE WATER
THICKENER OVERFLOW AND FILTRATE RETURN
GRIT CHAMBER
PRIMARY
CLARIFIERS
AERATION
BASINS
0/
SECONDARY
CLARIFIERS
f WASTE SLUDGE
MODEL
PLANT
GRAVITY
THICKENERS
VACUUM
FILTERS
T
LANDFILL
TO PISCATAWAY BAY
Figure 2. Schematic of the 5 mgd system of the
Piscataway plant secondary plant.
ANAEROBIC
DIGESTERS
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at 33% of total flow. At a flow of 5 mgd, the detention time in the reactors
is 4.5 hours. Air is supplied by diffusers located on the outside walls of
the reactors. The dissolved oxygen concentration in the reactors is control-
led by manually operated valves located above the water level. Normal oper-
ation called for control of the DO at approximately 2.5 mg/1 02- The dif-
fusers not only provide the required DO but also create a circular motion to
promote mixing. Surface sprays are located in each reactor along the center
feed channel.
The mixed liquor flows by gravity to the secondary settler inlet struc-
ture where the flow is split and fed to two settlers. At a flow of 2.5 mgd/
settler, the units provide 2.75 hours of detention time at a surface load-
ing of 650 gpd/sq ft. Centrifugal pumps return, the settled solids to the
reactors. Wasting of solids from the system is accomplished by diverting a
portion of the recycled flow to the gravity thickeners. The effluents from
the two clarifiers are combined and measured in a Parshall flume. The total
flow or a portion of the total flow is sent to the Model Plant's inlet struc-
ture. Excess flow not reporting to the Model Plant is diverted through a
Parshall flume to the polishing ponds which were installed at the plant as
an interim upgrading step pending completion of the 25 mgd system.
The waste solids from both the primary and secondary settlers flow by
gravity to the thickeners' inlet structure from which the flow is directed
to four gravity thickeners; two basins 35 ft in diameter and two basins
55 ft in diameter. The gravity thickening system was designed for a total
plant flow of 30 mgd. Since the total flow was less than 15 mgd during the
test period, only two or three of the thickeners were in operation.
The underflow from the thickeners was pumped to two anaerobic digesters.
The digesters were designed on a basis of 5 mgd wastewater flow but had to be
used at higher loadings because of a Prince Georges County Council ban on the
use of three fluidized bed solids incinerators. The WSSC was not allowed to
operate the incinerators because of possible air pollution. The capacity of
the digesters was the controlling factor in the solids handling system.
Because of the limited capacity of the digesters, thickener overflow, high
in suspended solids, was recycled to the inlet of the plant.
Following anaerobic digestion the underflow is vacuum filtered. The
cake is trucked to a farm for spreading. Because of health requirements,
only stabilized sludge can be vacuum filtered and disposed of on the land.
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IV. TERTIARY TREATMENT
Operation of the tertiary treatment plant consisted of treating the
secondary effluent from the 5 mgd step aeration system by lime clarification,
dual -media filtration and granular activated carbon adsorption. The solids
from the chemical clarification system were gravity thickened and dewatered
by solid bowl centrifuges . The dewatered cake was recalcined in a multiple
hearth furnace. The exhausted granular activated carbon was thermally
regenerated in another multiple hearth furnace.
Lime Treatment
In lime clarification of relatively low alkalinity wastewaters , two
options are available: two-stage high lime with intermediate recarbonation
or single-stage low lime. Since the alkalinity of the Piscataway wastewater
is less than 150 mg/1 as CaC03 and adequate flexibility was incorporated into
the design of the plant, both the single- and two-stage systems were evaluated.
A lime slurry, when added to a secondary wastewater, raises the pH of
the liquid to- produce chemical precipitation. Above pH 8.3, bicarbonate ions
are converted to carbonate ions which react with the available calcium ions
to precipitate calcium carbonate.
+ OH~ ^_ CaCO3 + H2O
Above neutral pH, the calcium ions react with phosphate ions in the presence
of hydroxyl ions to precipitate hydroxyapatite .
SCa** + 3H2PO~ + 70H~ ^ Ca5 (OH) (P04) 3 + 6H20
Millipore filtered samples from laboratory jar tests have shown that at pH 10
nearly all of the phosphate is precipitated, however, because of the lack of
a coagulant aid, the precipitate is not readily removed by simple settling.
By adding sufficient lime to increase the pH above 11.3, the magnesium ions ,
which are naturally present at between 3 and 6 mg/1 in Piscataway wastewater,
react to precipitate magnesium hydroxide, an excellent coagulant aid.
Mg++ + 2OH~ - ^_ Mg(OH)2
With the aid of the magnesium hydroxide, the precipitated phosphorus and
calcium carbonate coagulate and settle.
-------
When the pH of a low alkalinity wastewater has been increased to above
11.3 by the addition of calcium hydroxide, an excess of calcium ions exists.
The calcium ion concentration in Piscataway wastewater is increased from
approximately 30 mg/1 at neutral pH to approximately 100 mg/1 at pH 11.5.
The excess is produced because more calcium hydroxide is required to elevate
the pH to 11.5 than there are available phosphorus and bicarbonate ions to
react with the added calcium ions.
In the two-stage high lime system shown in Figure 3, lime slurry, waste-
water and recycled settled solids are mixed in the rapid mix zone of the
first clarifier to reach pH 11.5. The settled solids are recycled to in-
crease the rate of precipitation. Following rapid mix, flocculation and
settling, the effluent from the first clarifier flows to the recarbonation
basin where gaseous carbon dioxide is added for precipitation of the excess
calcium ions at pH 10.
Ca++ + 2OH~ + CO2 ;>.- CaC03 + H2O
Reducing the pH of the effluent from pH 11.5 directly to pH 7 by the addi-
tion of carbon dioxide resulted in solution of the calcium carbonate with
an increase in the hardness of the water and the loss of the potentially
recoverable calcium ions. The effluent containing the precipitated calcium
carbonate is then fed to a second clarifier. The calcium carbonate is
difficult to settle at pH 10 and a coagulant aid is required. Ferric
chloride, added at a dosage of 5 to 10 mg/1 as Fe+++, has been successfully
used as a coagulant aid.
Fe+++ + 30H~ ^^ Fe(OH)3
Following the settling of the calcium carbonate in the second clarifier,
carbon dioxide is added to the effluent to reduce the pH to 8 prior to
filtration. The neutralization of the effluent prevents calcium carbonate
scaling of the filter media.
OH~ + CO2 ^_ HCO~
CC>3 + C02 + H20 2HCO~
i
In the single stage lime system, shown in Figure 4, the pH of the waste-
water is increased to 10.5 by the addition of lime in the rapid mix zone of
the clarifier. As stated above, much of the phosphorus is precipitated along
with the available carbonate. However, without the precipitation of magnesium
hydroxide, good clarification does not occur and a coagulant aid is required.
Ferric chloride is added for this purpose to the rapid mix zone of the single
clarifier. The concentration of excess calcium ions in the effluent from the
single-stage system is approximately 50 to 60 mg/1 as Ca++ and recovery by
the addition of carbon dioxide is not economically feasible. Following settl-
ing in the single clarifier, the effluent is neutralized to pH 8 and filtered.
The single-stage lime system is not as effective as the two-stage for removing
phosphorus.
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SECONDARY EFFLUENT
FIRST REACTOR CLARIFIER
RECARBONATION BASIN
SECOND REACTOR CLARIFIER
FILTER INLET WELL
DUAL
MEDIA
FILTERS
STABILIZATION BASIN
CARBON ADSORBERS
POLISHING PONDS
CHLORINE CONTACT
PISCATAWAY BAY
Figure 3. Two-stage high lime tertiary process.
10
-------
SECONDARY EFFLUENT
REACTOR CLARIFIER
RECARBONATION BASIN
FILTER INLET WELL
DUAL
MEDIA
FILTERS
STABILIZATION BASIN
CARBON
ADSORBERS
POLISHING PONDS
\
CHLORINE CONTACT
I
PISCATAWAY BAY
Figure 4. Single stage low lime tertiary process.
11
-------
Neutralized effluent is applied to dual media gravity filters for the
removal of particulate materials including phosphorus and organics. The fil-
ter media consists of 18 inches of anthracite coal (effective size 0.85 to
0.95 mm) over 6 inches of fine sand (effective size 0.40 to 0.45 mm). The 24-
inch filter bed is supported by 12 inches of graded gravel (#10 mesh on top
to 1 inch on the bottom). The coarse to fine gradation of the media produces
in-depth filtration with the larger suspended particles being removed in the
anthracite coal and the smaller particles being removed in the fine sand.
The filters were normally operated at 3 gpm/sq ft and backwashed at a maxi-
mum of 20 gpm/sq ft, either on a predetermined time schedule or upon reaching
terminal head loss. Surface wash was applied during backwash to cleanse the
upper portion of the filter media to prevent the formation of mud balls. The
backwash water for both the filters and carbon columns is stored-and returned
to the head of the plant at a controlled rate.
The filter effluent is reduced to pH 7.5 in the water stabilization tank
by the addition of carbon dioxide. The reduction in pH is necessary to meet
stream discharge standards and to optimize activated carbon adsorption.
Activated Carbon Adsorption
The final unit process in the tertiary treatment scheme is activated car-
bon adsorption. The effluent from the stabilization tank is pumped at a rate
ojr.6.5 gpm/sq ft through a packed granular activated carbon bed which adsorbs
soluble organic materials from the wastewater. The adsorbed organics serve
as a food source for bacteria which multiply on the carbon to produce bio-
logical slimes. The biological activity, if controlled, can substantially
increase the life of the activated carbon. In order to control the activity
and to prevent excessive pressure losses through the packed carbon beds, back-
wash and surface wash of the carbon columns are necessary. The backwash rate
is 15 gpm/sq ft.
Carbon Regeneration
Following exhaustion of the activated carbon, the carbon is removed from
the column for thermal regeneration. It is transferred, at a controlled rate,
into a multiple hearth furnace where the regeneration takes place in 4 stages:
1. The wet carbon is dried by simple evaporation at temperatures
above 200°F.
2. Upon application of heat to the carbon grains at temperatures
above 600°F, the high molecular weight impurities on the carbon
will crack to produce gaseous hydrocarbons, hydrogen and water
vapor.
3. The final regeneration step is the gasification of the residue
from the pores of the carbon grains. This is accomplished
using steam (approximately 1 Ib steam/lb dry carbon) at
temperatures between 1700 and 1850°F. The gaseous products
of the reactions are carbon monoxide and hydrogen.
12
-------
4. The regenerated carbon is finally cooled rapidly to ambient
temperatures by water sprays.
Accurate control of the regeneration process is essential to maintain maximum
gasification of the organic residue without causing thermal destruction of
the granular carbon.
Lime Handling and Recovery
The chemical solids from the lime clarification system must be either
recovered for reuse or subjected to ultimate disposal. Laboratory tests
and material balance calculations show that at Piscataway in the single-
stage system, the solids production is approximately 2 lb/1000 gal with a
calcium carbonate concentration of 50%. The two-stage system will produce
approximately 4 lb/1000 gal with a 75% calcium carbonate concentration.
One of the objectives of this study was to determine the cost of solids
handling and calcination both with and without recovery of the calcium
carbonate.
For both land disposal or lime recovery, maximum dewatering of the
sludge is required to reduce operating costs,.i.e., trucking costs are based
on total weight and heat requirements increase with the quantity of water to
be evaporated in the furnace. In both the single stage and the two-stage
systems, the sludge is pumped to a gravity thickener. The sludge from the
second clarifier of the two-stage system is normally returned to the rapid
mix zone of the first stage to increase the calcium carbonate concentration,
thus improving settling. Wasting from the system occurs from the first
clarifier. Solids are removed from the clarifier at 3-5% total solids and
thickened to 10-15% by a gravity thickener. The overflow from the thickener
is returned to the first clarifier and the underflow solids are pumped to a
centrifuge for dewatering.
In dewatering the chemical sludges by centrifugation, two modes of
operation are available: total solids recovery and two stage wet classifi-
cation. The classification technique is used only when lime recalcination is
employed. In the total solids recovery mode, organic polymers are injected
into the centrifuge for maximum solids recovery. The centrate with a minimal
solids concentration is returned to the head of the plant. The wet cake may
be calcined for lime recovery or wasted. Thermal recalcination takes place
in a multiple hearth furnace by increasing the temperature of the cake to
1850°F with an auxiliary fuel source to convert the calcium carbonate to
calcium oxide.
CaCO3 ^- CaO + C02
The gaseous by-product, carbon dioxide, is recovered and used in the recarbon-
ation and stabilization tanks. The recalcined lime is slaked for reuse in
the lime clarification system.
13
-------
For the two stage wet classification mode, two centrifuges are operated
in series with the first machine being operated for capture of the calcium
carbonate in the centrifuge cake with the inerts reporting to the centrate.
At higher feed rates and without the addition of organic polymers, approxi-
mately 95% of the calcium carbonate in the sludge reports to the cake. The
centrate from the first centrifuge is pumped to the second machine where by
the addition of organic polymer the inert solids are captured in the cake.
The cake from the first machine is fed to the recalcination furnace and the
cake from the second machine is trucked to a landfill with the final centrate
being returned to the head of the plant. Since the cake from the first unit
is high in calcium carbonate, the recalcined lime contains a high percentage
of calcium oxide and wasting of the recalcined lime is not required. Addi-
tional lime, however, is required to maintain the high pH required in the
lime clarification processes.
In the total recovery operation, a portion of the recalcined lime, equal
to the inert build-up, is wasted from the system. Since calcium oxide is
also wasted with the inerts, it is necessary to add lime which is equivalent
to the amount wasted. In the two-stage system with secondary effluent, the
make-up rate is approximately 25%.
14
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V. DETAILED DESCRIPTION. OF THE MODEL TERTIARY PLANT FACILITY
The Model Plant equipment is presented in Tables 1-3. The flow rate
entering the plant is dependent on two operations. First, just past the
aerated grit chambers of the secondary treatment plant are located four
vertical constant speed centrifugal pumps each rated at 1,750 gpm at 5 ft
total dynamic head. The lift pumps were installed to maintain a constant
flow to the Model Plant. The second point of control is in the 30-inch
influent pipe to the Model Plant where an automatic flow control valve was
provided to control the influent flow rate to the Model Plant and isolate it
from process flow in case of an extended power failure or Model Plant
shutdown. A 14-inch magnetic flowmeter was provided to monitor the plant's
flow and transmit a signal to the main control panel to be recorded and
totalized. The flow signal is then transferred back to the automatic control
valve to maintain a predetermined flow as set on the controller at the main
panel. The inlet structure, two reactor clarifiers, recarbonation basin,
filter inlet tank and gravity thickener are located outside while the
remaining units are located inside the operations building.
Chemical Clarification
The flow of secondary effluent enters the influent structure which has
the capability of directing the flow to either of the two reactor clarifiers.
The flow pattern would depend on the preselected process to be evaluated
as shown in Figures 3 and 4. The effluent from the inlet structure flows
to the center draft tube of the first reactor clarifier, an 80 ft diameter
x 16 ft side wall depth tank. In the draft tube, lime slurry, ferric chloride
and/or polymer, settled solids and secondary effluent are combined under rapid
mixing conditions. The settled solids are drawn off the bottom of the
clarifier by the integral turbine mixer or may be recycled using an external
variable speed centrifugal pump'rated at 360 gpm.
The chemicals are applied by several methods. The lime handling system,
as shown in Figure 5, is controlled by an automatic non-clog ball valve
regulated by a signal sent to the control panel by a pH probe located in
the flocculation zone. The feed valve automatically opens and closes at
time intervals of one minute to prevent clogging. The 10% lime slurry is
recirculated at 120 gal/min from the lime slurry storage tank to the
clarifier through a series of 3 in steel lines and 2-1/2 in 100% dacron
firehose and returned to the storage tank. Most of the hose was installed
in an open trough to allow easy access for cleaning and replacement.
15
-------
TABLE 1
DESIGN DATA FOR MODEL PLANT EQUIPMENT
REACTOR CLARIFIER #1
1 unit - 80 ft dia x 16 ft SWD
Total detention time: 2.74 hr
Surface overflow rate: 1300 gpd/sq ft @ 5.6 mgd
RECARBONATION BASIN - 2 units
15.2 ft x 15 ft x 19.5 ft
Volume: 4446 cu ft or 33,256 gal
Detention Time: 9.6 min @ 5 mgd each
REACTOR CLARIFIER #2
1 unit - 70 ft dia x 16 ft SWD
Total detention time: 2.17 hr
Surface overflow rate: 1630 gpd/sq ft @ 5.4 mgd
INLET WELL
21 ft x 15 ft x 12 ft SWD
Volume: 3780 cu ft or 28,274 gal
Detention time: 7.6 min
GRAVITY FILTERS - 6 units
Surface Area: 242 sq ft each
Filter rate: 3.1 gpm/sq ft
Backwash rate: 20 gpm/sq ft
Media
Top Anthracite coal
Sand
Bottom Gravel
STABILIZATION BASIN - 1 unit
Volume: 34 ft x 17 ft x 12 ft SWD
6936 cu ft or 51,881 gal
Detention Time: 15 min
16
-------
TABLE 1 (Cont'd.)
DESIGN DATA FOR MODEL PLANT EQUIPMENT
ADSORBER INFLUENT PUMPS
3 pumps - (centrifugal)
50 hp 2000 gpm each
ACTIVATED CARBON COLUMN ADSORBERS - 6 units
Volume: 15 ft dia x 29 ft SWD: 5125 cu ft or 38,335 gal
Detention time in bed: 18.3 min each
Backwash rate: 20 gpm/sq ft
16 ft of granular activated carbon on a 1 ft gravel bed
GRAVITY THICKENER - 1 unit
50 ft dia x 10 ft SWD
Total Detention Time: 18.7 hr
Surface overflow rate: 115 gal/day sq ft
Solids loading: 31 Ib/day/sq ft
CENTRIFUGES - 2 units
Solid bowl super-D-canter
Size P-3400 30 hp Motor
Hydraulic capacity: 45 gpm
Scroll Speed: 3400 rpm
RECALCINATION FURNACE - 1 unit
Multiple hearth 6 levels
Design capacity total dry solids: 49,670 Ib/day
Dia: 19 ft
Total solids by Weight: 20%-55%
Moisture content by weight: 80%-45%
Maximum temperature in any hearth: 2000°F
Operating Pressures: negative 0.1 to negative 0.3 inches of water column
Four middle hearths have 2 burners each
Fuel Supply: #2 fuel oil
17
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TABLE 1 (Cont'd.)
DESIGN DATA FOR MODEL PLANT EQUIPMENT
CARBON REGENERATION FURNACE - 1 unit
Multiple hearth, 4 hearths plus afterburner section
Design Capacity: 4750 Ib/day dry carbon
Maximum Furnace temperature on any hearth: 1860°F
Afterburner Temperature Maximum: 1500°F
Fuel Consumption: Maximum 4,000 BTU/lb product
Fuel Supply: Natural gas for four burners and
#2 Fuel oil for two burners in afterburner chamber
18
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TABLE 2
MODEL PLANT PUMPS
PUMPS
1T Clarification System
2-Lime sludge recycling
2-Lime sludge wasting
2- Filter and carbon adsorption system
3-Adsorber pump
2-Backwash supplies
1-Filter decant (backwash effluent
•t^o headworks of Model Plant)
1-Adsorber decant (backwash effluent
to headworks of Model Plant)
SOLIDS HANDLING
1-Centrifuge feed
J.-Furnace feed
2-Classification pump
CHEMICAL FEED
' ' T'
4-polymer
4-Ferric Chloride
1-Ferric Chloride (transfer 40%
Fed3 to dilution tanks)
2-Lime slurry transfer •
2-Lime slurry recirculating
MISCELLANEOUS
^~ *" T T
2-ytility supply water
2-Operations building main sump pumps
centrifugal
progressive cavity
centrifugal
centrifugal
centrifugal
centrifugal
progressive cavity
progressive cavity
progressive cavity
positive displacement
positive displacement
centrifugal
centrifugal
centrifugal
centrifugal
centrifugal
360gpm
ISOgpm
2000gpm
SOOOgpm
200gpm
200gpm
55gpm
20gpm
45gpm
0.83gpm
0.83gpm
200gpm
40gpm
120gpm
350gpm
350gpm
19
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TABLE 3
MAJOR EQUIPMENT VENDORS
EQUIPMENT DESCRIPTION
Reactor-Clarifiers
Mixers
Dual Media Filters underdrains
Centrifuges
Lime Recalcination Furnace
Carbon Regeneration Furnace
Lime Slaking System
EQUIPMENT VENDOR
Rex Chainbelt
Mixing Equipment Co.
F.B. Leopold Co.
Sharpies
Nichols Corp.
Nichols Corp.
Wallace & Tiernan Inc.
20
-------
to
H
FROM
LIME
RECOVERY
FURNACE
GRAVIMETRIC
FEEDERS
SCREW
CONVEYORS
-FROM TRUCK
AUTOMATIC FEED CONTROL VALVE-
RECYCLE LINE-
LIME SLURRY
STORAGE TANK
'PUMP
PASTE SLAKER TRANSFER
SURGE
TANK
Figure 5. Lime handling system.
-------
The pebble virgin lime is delivered to the plant by truck in bulk and
stored in a 1240 cu ft storage tank. Both the pebble and recalcined lime
systems are equipped with gravimetric feeders to add a constant weight of
each material to the slaker. The lime is transported to the slaker via
screw conveyors.
The paste type lime slaker, rated at 2000 Ib/hr, converts the calcium
oxide to calcium hydroxide by the addition of water. The chemical reaction
is exothermic which increases rates of reaction. The paste is degritted
and diluted prior to discharge to the slurry tank. The lime slurry is pump-
ed from a 3 ft x 3 ft x 3 ft surge tank by a 200 gpm centrifugal pump to
the 12,000 gallon lime slurry storage tank. The storage tank is equipped
with a 15 hp mixer to maintain the solids in suspension.
In addition to the lime feed system, the plant is equipped with two
6,000 gallon ferric chloride storage tanks that receive 40% FeCl3 solution
via truck delivery. The 40% solution is pumped to one of the two 1500 gal-
lon tanks where the solution is diluted to 10% prior to use. Four manually
controlled diaphragm metering pumps, each rated at 50 gallons/hr, are pro-
vided to feed the ferric chloride. The polymer feed system consists of
two 600 gallon solution tanks and four manually controlled diaphragm feed
pumps each rated at 50 gallons/hr. The powdered polymer is received in
50 Ib bags and fed to the tanks by an operator. The units are equipped
with mechanical mixers. The dosages of these chemicals were varied as a
function of the supernatant clarity and settleability of the suspended
solids in the clarifier.
The reactor clarifiers, shown in Figure 6, consist of three major
sections: draft tube, flocculation zone and clarification zone. The draft
tube is located in the center section of the clarifier just below the tur-
bine mixer and is used as a rapid mix zone. The process water, chemicals
and recycled sludge are drawn by the mixer up and out into the flocculation
zone. The impeller provides rapid mix to disperse the chemicals and to
complete the chemical reactions and flocculation to promote particle growth.
The detention time of the flocculation zone is estimated as approximately
20 minutes. Another function of the clarifier is to promote the settling
of the precipitated solids. The process flow moves down inside the fiber-
glass cone separating the flocculation zone from the clarification zone.
This baffle cone extends about half way down into the tank and the process
flow moves through the flocculation zone and under the baffle. At this
point the liquid and solids begin to separate as the solids continue their
downward movement and the liquid begins to move upward. Testing of the
baffle showed that its length was insufficient to dissipate the kinetic
energy created by the mixer in the clarifier. The settled solids are col-
lected by a plow rake mechanism and moved to a center hopper. Sludge wast- !
ing and recycle exits are located in the center hopper of the unit. To '
prevent scale accumulation the sludge was transported through glass-lined
pipes from the hopper to the pumps. The clarified wastewater moves up
towards the surface and into a series of orifices in the pheripheral efflu-
ent trough. The orifices vary in size in an attempt to achieve a uniform
draw off from the clarifier. The total detention time in the first
22
-------
COLLECTOR DRIVE
TURBINE DRIVE
SUBMERGED ORIFICE
to
w
INFLUENT PIPE\
EFFLUENT TROUGH
CHEMICAL FEED LINES
POLYMER + FaCl3
^EFFLUENT PIPE
EXTERNAL SLUDGE
RECIRCULATION PIPE
SLUDGE COLLECTOR
SLUDGE DRAWOFF
Figure 6. Cross section of reactor clarifier.
-------
clarifier is 2.74 hr with an overflow rate of 1300 gpd/sq ft at a flow of
5.6 mgd.
From the first clarifier, the liquid flows through an open trough to
the recarbonation basin for pH adjustment by the addition of carbon dioxide.
The process flow enters the 32 ft x 15 ft x 12 ft tank and flows under and
through a redwood baffle wall into the first of two carbon dioxide mixing
zones. The first zone is used to accomplish approximately 75% of the re-
quired pH adjustment. A 30 hp turbine mixer and a 6-in flue gas feed line
are provided in the first zone. The effluent from the first stage passes
over a concrete baffle and into the equally sized second stage. The second
stage, which is primarily for final adjustment of the pH, includes a 25 hp
turbine mixer and a 4-in flue gas feed line. Both stages are equipped with
submerged pH probes that control the flow of carbon dioxide to the separate
stages by automatic gas regulating valves. The recorders and controllers
for the pH control loops are located on the main control panel.
A 3-in sludge recycle line was located in the first stage recarbonation
Zone. During the two-stage operation, settled solids were recycled at a rate
of 150 gpm from the center hopper of the second stage clarifier. The purpose
of the recycled solids was to promote the precipitation of calcium carbonate
and produce a more dense particle for settling. During the single-stage
operation, calcium carbonate recovery was not required and the recarbonation
basin was used to lower the pH below 8.3 where calcium carbonate is soluble.
Sprays were installed to reduce foaming produced by the large volumes of flue
gas introduced.
Gaseous carbon dioxide was provided from two sources; the recalcination
furnace and a 25-ton liquid carbon dioxide storage tank. The liquid carbon
dioxide source was used to supplement the carbon dioxide demand during
periods of zero feed to the furnace and was placed in full service when the
furnace was shut down. The concentration of carbon dioxide in the flue gas
from the recalcination furnace varied from 5% at operating temperatures and
zero feed to 12% at rated temperature and full feed capacity. At a 5% carbon
dioxide concentration and wastewater flow of 5 mgd, additional carbon dioxide
was required. The blending of two gases, at 5% CC>2 and 98% CO2, respectively,
produced pH control difficulties. The CO2 feed system with automatic pH
control had been designed to feed a 10% gaseous flow.
The effluent from the recarbonation basin flowed to the draft tube of
the second stage clarifier. The purpose of the second stage clarifier was
to capture the precipitated calcium carbonate produced during the two-stage
operation. During the single-stage operation the unit was not used. The
second clarifier is 70 ft in diameter with a side wall depth of 16 ft and
an overflow rate of 1,630 gpd/sq ft at 5.4 mgd. The design of the internal
mechanism and the chemical feed lines was identical to that for the first
clarifier. During the two-stage operation, ferric chloride was pumped from
solution tanks located in the operations building to the second-stage draft
tube to improve chemical clarification.
24
-------
Lime Recovery Operations
As stated earlier, the settled solids from the second stage clarifier
were recycled to the recarbonation basin by a centrifugal pump. Excess
solids from the second stage clarifier were wasted to the first stage clari-
fier via a variable speed progressive cavity pump. A similar pump was used
to waste solids from the first stage clarifier to the gravity thickener.
The thickener has an inside diameter of 50 ft and a side wall depth of 10 ft.
Thickening of the chemical sludge is enhanced by a mechanical rake operated
at 4.6 rph and powered by a 2 hp motor. The rake was designed to be raised
or lowered as a function of the solids concentration. The overflow from the
thickener flowed by gravity to the first stage clarifier. During normal
operation, the overflow from the thickener was relatively free of solids. At
the sludge production of the two-stage high lime system of 22,500 Ib/day, the
surface loading to the thickener was low at 11.5 Ib/day/sq ft.
During the operation of the furnace or sludge wasting from the total
system, the underflow from the gravity thickener was pumped to one of two
solid bowl centrifuges for dewatering. The centrifuges were located in the
operations building below the recalcination furnace. The centrifuges are
powered by 30 hp motors and operate at a scroll speed of 3400 rpm.
The sludge is fed by an inner tube to approximately the center of the
unit as shown in Fig. 7. As the sludge exits into the main drum area, the
solids are moved slowly by a screw conveyor to the cake end of the machine.
The liquid as it separates from the solids is mixed with polymer in order
to capture additional solids. The polymer is added to the centrifuge by
an additional inner tube just past the sludge addition point. The captured
solids, because of their increased weight due to the polymer addition, move
toward the cake end of the unit with the liquid moving toward the opposite
end and discharging from the machine as the centrate. The centrate flows
to the operations building's sump and is returned to the head of the model
plant.
The arrangement of the two centrifuges allowed two different modes of
operation of the solids dewatering system, total capture and two-stage wet
classification. The operation of the centrifuge as explained above and
shown in Fig. 8 provided total capture of the solids. The purpose is for
maximum capture of the total solids entering the centrifuge and produces a
centrate containing low solids. Classification for separation of the calcium
carbonate from the inert solids is accomplished using two centrifuges in
series as shown in Fig. 9. Thickener underflow is fed without polymer addi-
tion to the first centrifuge where most of the calcium carbonate reports to
the cake and is fed to the furnace. The centrate from the first unit con-
taining inerts and some calcium carbonate flows to a holding tank and is
pumped to the second centrifuge where by the addition of polymer the solids
are captured in the cake. The centrate from the second centrifuge returns
to the head of the plant via the operations building's sump pumps.
25
-------
HATE 0AM
(CENTRATE
DISCHARGE)
OUTSIDE COVER
PILLOW BLOCK
GEAR BOX
(O
TORQUE
MECHANISM
CONVEYOR OR SCROLL
POLYMER FEED PIPE
FEED NOZZLE /—CAKE DISCHARGE PORTS
1 flr
L_
r\
o
~Sr '.:•;... V
1
v^
/
SLUDGE FEED ZOI>
IE-
ftQfpfc
r •
II
i i i
a
\
POLYMER FEED ZONE
DRIVE PULLEY
HOLLOW FEED TUBE WITH POLYMER
ADDITION CAPABILITIES
POLYMER FEED POINT
-FEED SLUDGE
Figure 7. Solid bowl centrifuge section.
-------
CENTRIFUGE #2
-POLYMER
to
CENTRIFUGE #1
CENTRATE TO
HEAD OF PLANT
SLUDGE FEED
FROM THICKENER
•POLYMER
CAKE
PUMP TO FURNACE
OR TO LANDFILL
Figure 8. Centrifuge operation for total capture.
-------
SLUDGE FEED
CENTRIFUGE #1
POUMER
CENTRATE
CAKE
FEED TO
FURNACE
00
CENTRIFUGE #2
"1
CAKE
CENTRATE
TO HEAD
OF PLANT
TO SECONDARY
PLANT THICKENERS
Figure 9..Centrifuge operation for wet classification.
-------
During the classification operation, the cake from the second centri-
fuge, containing the inerts, was pumped to the secondary plant's thickeners.
The cakes from the parallel centrifuges during total capture and from the
first stage centrifuge during classification were carried by a conveyor
belt to a progressive cavity pump. This wet cake was either pumped to the
furnace for recalcination or drying, or pumped to a truck for land disposal.
An 18 ft diameter multiple hearth furnace was provided for recalci-
nation of the chemical solids. The furnace is comprised of 6 circular
hearths oriented horizontally and arranged for feeding solids at the top
level. Hearths 1, 2 and 6 have four rabble arms each while hearths 3,
4 and 5 are provided with two arms per hearth. The rabble arms are
equipped with teeth that are angled to rabble or rake the solids in a
spiral motion across the hearths. As seen in Fig 10, the solids are
directed on the first hearth from the outside of the furnace toward the
center shaft where the solids drop onto the second hearth. The procedure
is repeated through the six hearths with the solids reaching a maximum
temperature of 1850°F.
The moisture is evaporated from the lime sludge by hot combustion gases
which have been released from the auxiliary fuel burners mounted in hearths
2, 3, 4 and 5. The burners are supplied with #2 fuel oil that is mixed with
air supplied by a 20 hp combustion blower. The blowers are angled to produce
a circular motion of air and prevent the direct contact of the flame with the
furnace components. The pilot lights for each burner are fired by natural
gas. An automatic temperature control system is included.
The center shaft and rabble arms are cooled by a steady flow of air
supplied by a 10 hp fan motor. The exhaust gases are drawn off by a 100 hp
induced draft fan. The gases are first processed through a venturi scrubber
then pass into an entrainment separator. A flow of water of 164 gpm was
required by the venturi scrubber. A portion of the flue gases is compressed
to 8 psig and piped into the liquid treatment system for pH control.
The recalcined solids discharge from the sixth hearth to the thermal
disc product cooler at approximately 1600°F. The cooler includes three rows
of hollow water-cooled metal discs that rotate through the lime as it passes
through the cooler. The system is designed to cool the material from 1600°F
at a maximum capacity of 1520 Ib/hr. The cooling water is combined with
water flowing to the scrubber and is included in the total 164 gpm water
flow through the scrubber.
The cooled lime passes through the lump breaker and into an air-tight
rotary lock feeder which transports the material to the recalcination storage
bin by the introduction of pressurized air. The bin is equipped with an air-
powered mechanical hammer which prevents bridging. A bag type dust collec-
tion system is also included. The recalcined lime is fed to the gravimetric
feeder by a twin screw feeder.
29
-------
GAS OUT
LIME OUT^
RABBLE ARM
RABBLE TEETH
(a?^
Figure 10. Cross section of multiple hearth furnace.
30
-------
Filtration
The filter inlet well tank was located outside the operations building
and collected flow from the second stage clarifier during the two-stage
operation and from the recarbonation basin during the single stage operation.
The purpose of the tank was to provide level control for the gavity filters.
The inlet well was equipped with a 40 hp turbine mixer and a 6 in flue gas
line for final pH adjustment prior to filtration. The feed system was made
oversize as a safety factor to prevent calcium carbonate scaling of the dual
media filters if the primary recarbonation system failed to produce the
proper pH.
From the inlet well the flow passes into the operations building and is
fed to five of the six dual media filters. Each filter has a surface area
of 242 sq ft with a hydraulic loading of 3.1 gpm/sq ft at 1 mgd/filter.
A side view of a dual media filter is shown in Fig. 11. Normally, the sixth
filter is on standby and ready to be placed in service when required. The
filter media consists of the following:
Material Depth Size
Anthracite Coal 18 inches 0.85-0.95 mm
Sand 6 inches 0.40-0.45 mm
Support Gravel 3 Inches 10 mesh - 3/16 in
3 inches 3/16 in - 3/8 in
3 inches 3/8 in - 5/8 in
3 inches 5/8 in - 1 in
The flow through the filters is controlled by valves located at the
discharge of each unit. Signals from the main control panel network control
the five operating valves to maintain an even 5-way split of flow and a pre-
set water level. Once a control valve reaches 100% open and thus can no
longer control the flow, the filter-is taken out of service and backwashed.
However, for ease of scheduling, the units are backwashed every 24 hours at
approximately 8 ft of headless. A total of 12 ft of available head was
included in the design of the units.
The dual media filters are cleaned in 10 minutes by the combination of
surface wash and vertical backwash streams. Once the operator initiates the
backwash cycle the procedure continues automatically. The sequence includes
the following:
1. The units are drained to the drain troughs.
2. Backwash for 2 minutes at 10 gpm/sq ft.
31
-------
to
18" ANTHRACITE
12" SUPPORT GRAVEL
10" FILTER BLOCK
Ltrl
EL. 36.00'
1
&S
rtfk
v W.S.EL.33.00
4' t EL.30.50
TROUGH-^
— •
\.
1-7 TO T-12
— -
SURFACE
HGITATOR
I
— -~-
EL.19.50' /7\
tRETURN V
BACKWASH
rEL.18.001
t COLUMN- -
^
\
1
f-
S.18.751'^'
vy
1
y-El.33.001 1!| FLOOR
' ' „ j.EL.29.501 4" UTILITY WATER
-S^4vpStEL.27.42'— INFLUENT
-f^r70)tEl.24.92'_ DRAIN
t-R-^-LL^ FILTER
AFFLUENT
*? *"\ \ 8"BUTTERFLY VALVE
CYLINDER N (CV-12 TO CV-17)
FOR 14" 8"FLOW TUBE
B.F.V. (FE-12 TO FE-17)
•~ t COLUMN
BACKWASH SUPPLY
Figure 11. Cross section of dual media filter.
-------
3. Flow is increased to 20 gpm/sq ft for 6 minutes for
cleaning of the filter media.
4. The flow is decreased to 10 gpm/sq ft for 2 minutes
to allow uniform settling of the filter media.
5. The hydraulically driven surface wash mechanisms are
in service until the final 2 minutes.
The water supply for the backwash operation is filter effluent stored
in a 59,000 gallon tank (equivalent to 1.8 backwash) located below the
filter. Two backwash pumps are provided, each rated at 5000 gpm at 80 ft
of total dynamic head. The backwash water from the filter is collected in
another 59,000 gallon tank and returned to the head of the plant at the
inlet structure at a rate of 250 gpm. The reduced flow is provided to pre-
vent shock loads, both hydraulic and solids, to the first clarifier. The
water for the surface wash mechanism is carbon adsorber effluent.
The effluents from the operating filters are collected in a common line
and pass into the stabilization basin where flue gas is used to further ad-
just the pH, if necessary, prior to carbon adsorption. The 34 ft x 17 ft x
12 ft tank is covered and vented to prevent discharge of the toxic flue gas
within the operations building. The unit also provides flow equalization
ahead of the carbon adsorbers. Two 15 hp turbine mixers with two 4 in flue
gas lines are installed. The effluent from the stabilization basin passes
through over-under baffles and into the carbon adsorption wet well.
Carbon Adsorption
The carbon adsorption system consists of six downflow pressurized
vessels arranged in sets of two to provide three parallel trains of two col-
ums each. At a flow of 5 mgd, the loading to each column is 6.5 gpm/sq ft.
A depth of 16 ft of activated carbon provides 18 minutes of Empty Bed Con-
tact Time (EBCT) per column. The granular activated carbon was Filtrasorb
300* (8 x 30 mesh). The gravel support media for the activated carbon was
similar to that provided in the dual media filters. Each of the three car-
bon trains is provided with a 50 hp, 2000 gpm centrifugal pump. The present
flow control loop used to split and control the flow through the carbon ad-
sorption system is similar to the system provided in the filters. The level
in the carbon column inlet well was maintained by 3 control valves placed at
the ends of the 3 parallel carbon trains. The control valves compensate for
changing headlosses and maintain an even 3-way split of the flow and steady
level in the inlet well by an analog logic network located in the main con-
trol panel.
The lead carbon column of each train was backwashed daily with the
final column being backwashed every other day. The backwash cycle is initi-
ated by the operator but then proceeds automatically. A series of on-off
valves automatically isolate the columns to be backwashed, thus during
*A product of Calgon Corp., Pittsburgh, PA.
33
-------
backwash, for approximately 10 minutes, one-third of the flow receives the
treatment of only one carbon column. The time intervals of backwash include
2 minutes at 8.5 gpm/sq ft, 6 minutes at 17.5 gpm/sq ft and 2 minutes at 8.5
gpm/sq ft. A hydraulic surface wash mechanism is located 6 inches above the
top of the carbon bed. The filters and carbon columns utilize the common
59,000 gallon backwash water supply tank. The backwash water flows through
the column into the carbon adsorption sump tank from which it is returned to
the head of the Model Plant at a rate of 250 gpm.
The carbon regeneration system is shown in Fig. 12. The exhausted car-
bon is removed from the activated carbon column through the four funnels
located in the underdrain of each column (Fig. 13) and hydraulically carried
to the exhausted carbon storage tank. The storage tank is equipped with a
1 in eductor to feed the carbon to the dewatering screw at a rate of 4750
Ib/day. Initially, the transfer line was 1 in fiberglass with 90° elbows
but was later converted to 1-1/4 in flexible tubing to prevent plugging.
The concentration of carbon in the slurry is approximately 1 Ib/gallon. The
carbon slurry is dewatered in an inclined screw conveyor.
The dewatered carbon is fed to the second hearth of the 54-in diameter
multiple hearth furnace. The carbon furnace, like the lime recalcination
furnace, is a multiple hearth. The unit has five hearths with rabble arms on
four levels to move the carbon downward through the hearths. The top hearth
is used as an afterburner for the ignition of the off gases from the carbon
regeneration process. The exhaust gases pass through a venturi scrubber and
entrainment separator. The wash water from the separator is returned to the
model plant's sump.
The carbon passes through the four lower hearths and discharges to the
quench tank where the carbon is rapidly cooled in water. Two of the four
hearths which receive carbon are equipped with burners which are fired by
natural gas. A blower supplies a steady flow of air to cool the center
shaft and rabble arms. From the quench tank, the activated carbon is
hydraulically transported to the regenerated carbon storage tank where it
is held until the regeneration of another column is required.
The effluents from the three carbon systems are collected in a common
line and flow to the polishing ponds. The effluent from the ponds receiving
carbon effluents is combined with the effluents from the ponds containing
Piscataway secondary effluent, chlorinated and discharged to Piscataway
Bay.
-------
CARBON ADSORBERS
t
i
1
1
1
1
SPENT CARBON STORAGE TANK
J/f
MULTIPLE HEARTH REGENERATION FURNACE
\7
QUENCH TANK
REGENERATED CARBON STORAGE TANK
Fjgure 12. Flow schematic for carbon regeneration.
35
-------
-ISft.O.D.-
GRANULAR CARBON
FILTER GRAVEL
PROCESS EFFLUENT
4-Carbon draw-off hoppers used to transfer carbon
from the adsorbers to the spent storage tank.
Figure 13. Cross section of carbon adsorber underdrain.
36
-------
VI. RESULTS OF TWO-STAGE HIGH LIME EVALUATION
The operation of the two-stage high lime system was sustained at the
design flow rate for 36 days in October and November, 1973, a period between
two failures of the reactor clarifiers.
Secondary Operation
The operation of the step aeration activated sludge system described in
Table 4 corresponds to the 36 days of operation of the tertiary facility
and was typical of that achieved during the two years of the grant period.
The mixed liquor suspended solids concentration of 2133 mg/1 is an average of
samples taken at the quarter points along the reactor. As seen by the aver-
age decrease in alkalinity from 135 to 97 mg/1 as CaCO^, the activated sludge
system was nitrifying. The reduction in alkalinity is a result of the pro-
duction of nitric acid during nitrification. The variations in alkalinity
and TKN in the secondary effluent are shown in Fig. 14. The degree of nitri-
fication was inconsistent, most probably because of the irregular wasting
schedule from the secondary settlers. The SRT of 6.6 days is based on the
average wasting rate for the 36 day operating period. The average concen-
trations of BOD and suspended solids for the raw wastewater, primary effluent
and secondary effluent are presented in Tables 5 and 6. The effect of the
recycle of solids to the grit chamber is reflected in the high values for
these parameters in the primary effluent.
The phosphorus concentrations are given in Table 7. During the evalu-
ation period for the tertiary lime treatment systems, 80 mg/1 alum was being
added to the aeration tanks of the 25 mgd activated sludge system to improve
removals of phosphorus and suspended solids. Return streams to the head of
the plant containing high concentrations of alum sludge and precipitated
phosphate entered the 5 mgd secondary system and produced high phosphorus
concentrations in the primary effluent. The overall effect of the return
stream was an increase in phosphorus removal through the old 5 mgd secondary
plant. During the 36 days of operation, the secondary plant removed 55.7%
of the incoming phosphorus. The average nitrogen concentrations are given
in Table 8, but do not reflect the variability as seen in Fig. 14.
Tertiary Treatment
During the 36 days of continuous operation the influent flow to the Model
Tertiary Plant averaged 4.586 mgd. The hydraulic loadings to the unit proc-
esses based on influent flow plus recycle are presented in Table 9. The
recycled plant water was 11.3% of the total flow. A summary of the plant
recycle flows is presented in Table 10.
37
-------
TABLE 4
OPERATING CONDITIONS OF THE PISCATAWAY SECONDARY PLANT
DURING THE HIGH LIME PROCESS EVALUATION
DAILY FLOW, mgd 5.675
DETENTION TIME, hr 4.4
MLSS, mg/1 2133
RECYCLE RATE, % 37
SOLIDS IN RECYCLE, mg/1 7970
WASTE RATE, 1000 gal/day 38.1
SVI, ml/gin 96
SRT, days 6.6
F/M, Ib BOD5/lb MLVSS 0.41
RAW WASTEWATER pH 7.2
SECONDARY EFFLUENT pH 7.4
RAW WASTEWATER TEMPERATURE, °F 65
RAW WASTEWATER ALKALINITY, mg/1 CaCO 135
SECONDARY EFFLUENT ALKALINITY, mg/1 CaCO. 97
TABLE 5
REMOVAL OF BIOCHEMICAL OXYGEN DEMAND (BOD 5 DAY)
DURING THE HIGH LIME PROCESS EVALUATION
mg/1 % Removal
Raw 141.0
Primary 145.0
Secondary 16.5 88.3
Lime clarified 5.9 95.8
Filtered 5.7 96.0
Carbon Adsorption 4.0 97.2
Nots *
"Secondary Plant recycle enters between the raw sample point and
the primary clarifier.
38
-------
140
CO
O
u
0
u
100
10
O)
E
60
ALKALINITY
12
3
(D
Q
(/i
z
16
24
32
DAYS
Figure 14. Comparison of alkalinity and TKN of secondary effluent.
-------
TABLE 6
REMOVAL OF SUSPENDED SOLIDS DURING EVALUATION
OF THE HIGH LIME PROCESS
mg/1 % Removal
Raw 121
Primary 183
Secondary 27.5 77.3
Lime Clarified 21 82.6
Filtered 6 95.0
Carbon Adsorption 2.5 97.9
Note:
Secondary Plant recycle enters between the raw sample .point
and the primary clarifier.
TABLE 7
REMOVAL OF TOTAL PHOSPHORUS (AS P) DURING EVALUATION
OF THE HIGH LIME PROCESS
mg/1 % Removal
Raw 7.90
Primary 9.60
Secondary 3.50 55.7
Lime Clarified 0.26 96.7
Filtered 0.20 97.5
Carbon Adsorption 0.10 98.7
Note:
Secondary Plant recycle enters between the raw sample point
and the primary clarifier
40
-------
TABLE 8
REMOVAL OF NITROGEN COMPOUNDS DURING EVALUATION
OF THE HIGH LIME PROCESS
Raw
Primary
Secondary
Lime. Clarified
Filtered
Carbon Adsorption
*A11 values mg/1 as N
NH
13.1
TKN
16.2
NO
1.1
0.1
TOTAL N
17.4
4.5
5.4
4.1
2.9
5.2
6.0
3.3
6.7
6.4
5.9
8.1
.4
.3
.1
12.3
12.7
11.5
TABLE 9
LOADING RATES DURING EVALUATION OF THE HIGH LIME PROCESS
Secondary Plant
Flow 5.7 mgd
Primary Clarifiers
Secondary Clarifiers
2-60 ft dia.
2-70 ft dia.
Model AWT Plant
External Flow and Recycle = Total Flow
4.586 mgd + 0.588 mgd = 5.174 mgd
First Stage Reactor Clarifier 1-80 ft dia.
Second Stage Reactor Clarifier 1-70 ft dia.
Dual Media Filtration (5 units) 242 sq ft/unit
Carbon Adsorption
Column Set #1 Avg. Flow 2.063 mgd
Column Set #2 Avg. Flow 1.420 mgd
Column Set #3 Avg. Flow 1.309 mgd
All loadings are based on an average flow of 5.174 mgd.
100R gpd/sq ft
74-1 gpd/sq ft
1029 gpd/sq ft
1344 gpd/sq ft
4276 gpd/sq ft
or
2.97 gpm/sq ft
8.1 gpm/sq ft
5.6 gpm/sq ft
5.1 gpm/sq ft
41
-------
TABLE 10
PLANT RECYCLE FLOWS DURING HIGH LIME EVALUATION
SOURCES TOTAL GALLONS/DAY % OF FLOW
1 FILTER BACKWASH
38,720 gal x 6 filters/day 232,320 4.49
2 CARBON COLUMN BACKWASH
24,017 gal x 4 columns/day 96,068 1.86
3 RECALCINATION FURNACE
164 gal/min x 1440 min/day 236,160 4.56
4 Misc. - (centrate, pump sealing
Water, flushing & wash water) 23,452 0.45
TOTAL 588,000 11.36
All percentages are based on an average flow of 5.174 mgd.
42
-------
The first reactor clarifier was maintained at pH 11.45 by an average
lime dose of 257 mg/1 as CaO. The lime dose calculation is based on total
pounds of recalcined and virgin pebble limes added to the system with aver-
age Available Lime Indexes (ALI) of 60 and 87%, respectively. The quanti-
ties of chemicals added to the system during the 36 days of operation are
presented in Table 11. During the operation of the two stage system, the
solids handling system was operated for total capture with wasting of the
recalcined lime to prevent build-up of inerts in the system. As a result,
the low ALI of 60% was produced. The recalcined lime accounted for 75% of
the total material added to the system, or 68% of the required calcium oxide.
The effluent from the first clarifier was reduced to an average pH 10
in the recarbonation basin. During periods of zero or low sludge feed to
the furnace, the percent of carbon dioxide in the flue gas decreased to
approximately 5% which was insufficient to maintain the proper pH. With
adequate sludge feed to the furnace the carbon dioxide concentration was
approximately 12% which was sufficient for pH control. Additional carbon
dipx^de feed from the liquid storage tank was required on 14 of the 36
operating days. Because of the difficulty in maintaining proper control of
the 98% and 5% carbon dioxide feed systems, an accurate measurement of the
amounts of carbon dioxide added to the system, was not possible. Sludge
from the second clarifier was recycled to the recarbonation basin at a rate
of 5% of the average daily flow.
Following recarbonation, the effluent was fed to the second stage
clarifier for settling of the precipitated calcium carbonate. A ferric
chloride solution at an average concentration of 17.8 mg/1 as FeCl3 was
added to the draft tube of the second unit to improve clarification. The
waste sludge from the second clarifier was pumped to the first clarifier1s
draft tube at a rate of 0.66% of average daily flow. Total solids wasting
from the system was accomplished by pumping the sludge from the first clari-
fier to the gravity thickener at an average pumping rate of 0.7% of average
flow.
The results of the operation of the chemical clarification system are
presented in Tables 5 through 8 and in Table 12. As expected, the lime
clarification system reduced the phosphorus concentration to low levels,
0.26 mg/1 as P. With efficient.insolubilization of the BOD in the second-
ary treatment system and capture of the secondary effluent suspended solids
in the chemical clarification system, an average BOD concentration of 5.9
mg/1 was produced from the lime treatment system.
The second stage effluent was reduced to pH 8 by the addition of flue
gas and supplemental carbon dioxide in the filter inlet well prior to dual
media filtration. As with the recarbonation basin, an accurate measurement
of the amount of carbon dioxide could not be determined.
Five dual media filters were used continuously with another filter
either in backwash or standby. The flow rate of the filters averaged 2.97
gpm/sq f£. The filters were backwashed on a 24-hour cycle for ease of
operation at an average head loss of 8 ft.
43
-------
TABLE 11
CHEMICAL USAGE IN THE HIGH LIME PROCESS
Lime
Total Virgin Pounds 137,079
Average Available Lime Index (ALI), % 87
Available CaO, Ib 119,259
Average Daily Usage, Ib • 3,138
Average Daily Dose, mg/1 87
Total Recalcined Pounds 426,618
Average Available Lime Index (ALI),% 60
Available CaO, Ib 255,971
Average Daily Usage, Ib 6,731
Average Daily Dose, mg/1 170
Ferric Chloride (FeCl3)
Total Pounds added to clarifier,lb 23,496
Average Daily Usage,Ib 618
Average Daily Dose, mg/1 17.8
Polymer Usage (Centrifuge Only)
Total Pounds used 568.3
Average Daily Usage, Ib 15.0
Pounds Polymer/Ton of Dry Sludge 0.25
44
-------
TABLE 12
REMOVALS OF CHEMICAL OXYGEN DEMAND (COD) AND
TOTAL ORGANIC CARBON (TOC) DURING EVALUATION
OF THE HIGH LIME PROCESS
COD TOC
ng/1 mg/1 as C
Raw
Primary
Secondary 34.8 12.3
Clarified 13.4
Filtered 24.7 7.5
Carbon Adsorption 13.4 1.8
The dual media filters were not effective in removing additional
materials with the exception of suspended solids. The calculated removal
of suspended solids in the filters was 15 mg/1. However, since the suspend-
ed solids in the influent to the filter system consists mainly of precipi-
tated calcium carbonate, a portion of which is solubilized in the filter
inlet well, an accurate efficiency of the filters alone cannot be deter-
mined. However, if one assumes that 15 mg/1 of suspended solids is captured
in one 24-hour run at 1 mgd, an estimate of the efficiency of the filters
can be determined. A total of 125 Ib of solids would be captured in each
filter for a loading of 0.52 Ib/sq ft/cycle. By using a termination head-
loss of 8 ft, then the loading can be expressed as 0.065 Ib/sq ft/ft of
headless. These loadings are quite reasonable considering that the units
contain 24 inches of filter media compared to 36 inches of media in most
designs.
The filter effluent was reduced to pH 7.2 by the addition of flue gas
in the stabilization basin and pumped to the three parallel activated car-
bon adsorption systems. As seen in Table 9, the hydraulic loadings to the
three systems averaged 8.1, 5.6 and 5.1 gpm/sq ft. The inconsistency in
the loadings was due to the inability of the control system to maintain an
even split in flow. The lead column of each system was backwashed every
24 hours with the final column being backwashed every 48 hours.
45
-------
Approximately 24,000 gallons of water was required to backwash a carbon col-
umn as compared to 39,000 gallons to backwash a dual media filter. During the
backwash of a column, treatment by only one column in that train was provided.
The performance of the activated carbon system based on composite
samples from the three trains is presented in Tables 5 through 8. The
results are the averages for the 36 days of operation. As seen in Table 12,
the units removed approximately 75% of the inlet TOC and 50% of the COD
which is typical of activated carbon operation at other locations. The BOD
analysis was uninhibited and included the effects of sample nitrification
in addition to oxidation of organics. The degree of nitrification in the
activated carbon columns, from 5.9 to 8.1 mg/1 of NO^-N as shown in Table 8,
was typical of the operation of the units throughout the study.
The performances of the individual columns during the 36-day high lime
evaluation are presented in Tables 13, 14 and 15. The cumulative loadings
to the carbon columns, based on the cumulative flow through each column
from March 1973 through November 1973 are presented in Table 16. Column
T-16 was shut down for nearly two months prior to October 13, 1973, because
of a broken seat in the discharge valve during which period only T-17 was
in operation. At reduced flows, only columns T-14 and 15 were in operation
which accounts for their higher loadings and cumulative flows. BOD loadings
are not presented because of the effect of nitrification on the BOD analysis.
With the failure of the chemical clarifiers on November 18, 1973, carbon
column T-14 was prepared for regeneration. As seen in Tables 13 and 16,
the carbon was not exhausted but was still efficient in removing organics.
It was obvious, however, that the plant was to be down for an extended period
and the information and experience obtained from a regeneration cycle within
the grant period was considered to be important. In addition, with the re-
mainder of the plant shut down, full operator attention could be focused on
the carbon regeneration system. The results of the carbon regeneration
operation are presented in detail in Section VII.
The waste solids from the first clarifier, which included the solids
wasted from the second clarifier, were pumped to the gravity thickener for
solids processing. The solids concentration in the underflow from the first
clarifier generally ranged from 7 to 10% and varied as a function of the
pumping rate, which during the 36 days of operation averaged 0.70% of the
influent flow or 32,000 gallons/day.
Three methods were used to determine the material balances around the
clarification and solids handling system using the following data:
1. Chemical analyses of the liquid streams and measured daily
influent flow rates.
2. Chemical analyses of the sludge streams and measured sludge flows.
3. Total solids concentrations and measured sludge flows.
46
-------
TABLE 13
PERFORMANCE OF CARBON ADSORBER TRAIN #1 DURING
EVALUATION OF THE HIGH LIME PROCESS
Columns T-14 and T-15
Average Flow 2.063 mgd
Cumulative flow from March 1973 - November 1973 199.907 mil gal
Total Organic Carbon (TOC) mg/1 as C % Removed
Influent 7.5
Intermediate 3.2 57.3
Effluent 2.3 69.3
Biochemical Oxygen Demand (BOD) mg/1
Influent 6.2
Intermediate 5.8 6.4
Effluent 4.0 35.5
Chemical Oxygen Demand (COD) mg/1
Influent 24.9
Intermediate 17.4 30.1
Effluent 15.1 39.4
Organic concentrations based on data from October 14 to November 18, 1973.
47
-------
TABLE 14
PERFORMANCE OF CARBON ADSORBER TRAIN #2 DURING
EVALUATION OF THE HIGH LIME PROCESS
Columns T-16 and T-17
Average Flow 1.420 mgd
Cumulative flow from March 1973 - November 1973 173.151 mil gal
Total Organic Carbon (TOC) mg/1 as C % Removed
Influent 6.5
Intermediate
Effluent 2.4 63.1
Biochemical Oxygen Demand (BOD) mg/1
Influent 6.0
Intermediate
Effluent 4.6 23.3
Chemical Oxygen Demand (COD) mg/1
Influent 24.2 —
Intermediate
Effluent 13.7 43.4
Organic concentrations based on data from October 14 to November 18, 1973.
48
-------
TABLE 15
PERFORMANCE OF CARBON ADSORBER TRAIN #3 DURING
EVALUATION OF THE HIGH LIME PROCESS
Columns T-18 and T-19
Average Flow 1.309 mgd
Cumulative flow from March 1973 - November 1973 126.396 mil gal
Total Organic Carbon (TOC) rr.g/1 as C % Removed
Influent 7.6
Intermediate 2.6 65.8
Effluent 1.5 80.3
Biochemical Oxygen Demand (BOD) mg/1
Influent 6.3
Intermediate 5.6 11.1
Effluent 4.5 28.6
Chemical Oxygen Demand (COD) mg/1
Influent 25.4
Intermediate 15.4 39.4
Effluent 12.4 51.2
i
Organic concentration based on data from October 14 to November 18, 1973.
49
-------
TABLE 16
CUMULATIVE COD AND TOG LOADINGS ON CARBON AT THE
END OF THE 36-DAY HIGH LIME EVALUATION
Ib COD/lb Carbon Cumulative Ib TOC/lb Carbon
Flow, mil gal
Set 1
T-14 0.21155 199.907 0.09613
T-15 0.06893 199.907 0.01812
Set 2
T-16 0.09855 132.801 0.05230
T-17 0.09553 173.151 0.03419
Set 3
T-18 0.12065 126.396 0.06207
T-19 0.04086 126.396 0.01222
TABLE 17
SOLIDS MATERIAL BALANCES FOR THE HIGH LIME EVALUATION
FIRST CLARIFIER
Solids Captured 17,400 Ib/day
Solids Wasted to Thickener 22,400 Ib/day
SECOND CLARIFIER
Solids Captured 5,100 Ib/day
Solids Wasted to First Clarifier 11,800 Ib/day
TOTAL SOLIDS REPORTING TO FIRST CLARIFIER 22,500 Ib/day
SOLIDS REPORTING TO CENTRIFUGE 24,700 Ib/day
*SOLIDS REPORTING TO FURNACE 20,200 Ib/day
*Measured by gravimetric feeder following recalcination furnace
assuming a 75% loss through furnace.
50
-------
The accuracy of method two is limited because of the analytical and sampling
errors associated with the highly concentrated sludge streams. Method three
was based on average solids concentrations which in practice were quite vari-
able and therefore difficult to sample. The data for the liquid streams were
considered to be the most reliable because of the accuracy of the chemical
analyses in combination with good flow measurements.
Accepting these limitations, material balances around the clarifiers
were calculated and are presented in Table 17. The average solids captured
in the first and second stage clarifiers of 17,400 and 5,100 Ib/day, respec-
tively, are based on chemical analyses and flow measurements of the liquid
streams. The inert solids in the recalcined lime and return flow (filter
and carbon column backwash, centrate and scrubber) water are included in
the solids captured in the first clarifier. The solids wasted to the first
clarifier and thickener are based on average daily waste rates and total
solids concentrations. The amount which appears to be in gross error is the
11,800 Ib/day that reported from the second clarifier to the first clarifier.
The concentration of total solids in the waste sludge was variable and no
doubt resulted in an erroneous average number. By using the liquid stream
approach then, as seen in Table 17, the amount of solids (22,500) that re-
ported to the clarifier is close to the amount of solids that was pumped
from the first clarifier to the thickener (22,400). The amount that report-
ed to the thickener was based on measured total solids and waste flow rates.
The data show that the combination of clarifier-thickener was in solids
balance.
During the operation of the recalcination furnace, the underflow from
the gravity thickener, containing approximately 20% total solids, was
pumped to one of the two solid bowl centrifuges for dewatering. A moder-
ately anionic polymer was added to the centrifuge at a rate of 0.61 Ib
polymer/ton dry solids for total capture. Sludge cake with a suspended
solids concentration of 38.5% was produced and pumped to the top of the
recalcination furnace.
The combination of centrifugation and recalcination was not operated
pontinuously during the 36 days of high lime evaluation. The total solids
production of 22,500 Ib/day in the clarifiers was less than the capacity of
the furnace. Since the solids handling system required high manpower, it
was operated only to produce sufficient recalcined lime and/or maintain the
sludge blanket level in the gravity thickener.
A total of 545,818 Ib of recalcined lime was produced at an average
available lime index of 60%. Since the centrifuge was operated for total
capture of the solids, wasting of the recalcined lime to a sanitary landfill
was required to prevent the buildup of inerts in the lime. Based on pre-
liminary calculations, an estimated waste rate of 25% was established. A
total of 144,200 Ib of recalcined lime, or 25% of the total production, was
wasted. The recalcined and pebble limes were mixed at a 75/25% ratio by
twg gravimetric feeders prior to slaking. With the exception of increased
grit production, the paste slaker operation was reasonably successful.
51
-------
A heat balance of the recalcination furnace was calculated based on
average concentrations of materials in the centrifuge cake and average oper-
ating temperatures. The results of the heat balance are presented in the
Appendix. The feed to the furnace averaged 3315 lb;of wet sludge/hr or
1276 dry Ib/hr.
In order to determine if the lime recovery system could supply on a
daily basis, the recalcined lime actually added to the system, the furnace
was assumed to be operated at 1276 dry Ib/hr until the daily sludge pro-
duction was recalcined. The length of time for furnace operations, based
on a daily sludge production of 22,500 dry Ib/day, was determined to be
17.6 hr. The total amount of CaO produced in the furnace, based on the
data in the Appendix, would equal 9046 Ib/day (514.6 Ib/hr x 17.6 hr).
Assuming a waste rate of 26% or 2351 Ib/day, the total calcium oxide avail-
able for reuse would equal 6695 Ib/day. As seen in Table 11, the average
daily usage of CaO was actually 6731 Ib/day. Based on this information it
can be assumed that the clarification and recalcination systems were in
reasonable balance.
The heat balance around the furnace was good with 94% of the fuel used
accounted for in the calculation. The importance of reducing the moisture
concentration of the sludge feed should be noted. The BTU requirement for
the evaporation of the moisture accounted for 58% of the total BTU input.
52
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VII. RESULTS OF SINGLE-STAGE LOW LIME EVALUATION
The second major operational system to be tested was single-stage low
lime. This system was started on April 10, 1974, following replacement of
main gears in both clarifiers, and operated for 89 days until July 8, 1974,
when the induced draft fan on the recalcination furnace malfunctioned.
Secondary Operation
The operation of the secondary facility was generally typical of that
described for the two-stage high lime mode. Several factors were different,
however, as shown in Table 18 which represents the operational results for
the 89-day operating period. The mixed liquor suspended solids concentra-
tion of 2600 mg/1 is an average of four samples taken at the quarter points
along the reactor. The major differences in the secondary operations for
the high lime and low lime operations were in the sludge retention times
and the recycle rates. As seen by the decrease in alkalinity from 145 to
119 mg/1, the activated sludge system was nitrifying. The waste rate of
50,280 gal/day was the average rate for the total period. Because of the
configurations of the secondary facility and the two separate systems used,
wasting from the aeration basins was conducted as a batch operation.
The average concentrations of BOD and suspended solids for the raw
wastewater, primary effluent and secondary effluent are presented in Tables
19 and 20. The effects of the recycle of solids to the grit chamber are
reflected in the high values for the primary effluent. The recycled solids
to the grit chamber are from the overloaded thickeners and solids handling
system. The phosphorus concentrations, as given in Table 21, again show
the results of recycling the aluminum phosphate enriched flow to the grit
chambers. A secondary plant removal of 58.5% of the raw sewage phosphorus
was obtained. The nitrogen concentrations are given in Table 22.
Tertiary Treatment
The single stage low lime tertiary system as shown in Fig. 4, operated
at an average daily flow of 4.092 mgd. The hydraulic loadings to the unit
processes based on influent flows plus recycle are presented in Table 23.
As shown in Table 24, the recycle of the plant water accounted for 12% of
the total flow in the tertiary system.
The reactor clarifier was maintained at pH 10.4 by an average lime dose
of 113.4 mg/1 and a ferric chloride dose of 25.2 mg/1 as FeCl3 as shown in
Table 25. All solids formed in the lime clarification system were precipi-
tated in the reactor-clarifier and were pumped from the center collection
hoppsr to the gravity thickener. As seen in Table 25, most of the lime used
53
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TABLE 18
OPERATING CONDITIONS OF THE PISCATAWAY SECONDARY
OPERATION DURING THE LOW LIME PROCESS EVALUATION
Daily Flow, mgd ' 5.276
Aeration Tank Detention Time, hr 3.0
MLSS (% Volatile), mg/1 2600 (64)
Recycle Rate, % 48
Solids in Recycle, mg/1 6922
Waste Rate, 1000 gal/day 50.28
SVI, ml/gm 104
SRT, days 2.67
F/M, Ib BOD5/lb MLVSS 0.47
Raw Wastewater, pH 7.0
Secondary Effluent, pH 7.3
Raw Wastewater Temperature, °F 62
Raw Wastewater Alkalinity, mg/1 CaC03 145
Secondary Effluent Alkalinity, mg/1 CaCOs 119
TABLE 19
REMOVAL OF BIOCHEMICAL OXYGEN DEMAND (BOD 5 DAY)
DURING EVALUATION OF THE LOW LIME PROCESS
% Removal
Raw
Primary
Secondary 16.0 87.1
Clarified 8.2 93.4
Filtered 6.6 94.7
Carbon Adsorption 2.3 98.2
Note:
Secondary Plant recycle enters between the raw sample
point and the primary clarifier.
54
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TABLE 20
REMOVAL OF SUSPENDED SOLIDS DURING EVALUATION
OF THE LOW LIME PROCESS
% Removal
Raw 130.2
Primary 376.8
Secondary 18.5 85.8
Lime Clarified 15.8 87.9
Filtered 10.2 92.2
Carbon Adsorption 3.5 97.3
Note:
Secondary Plant recycle enters between the raw sample
point and the primary clarifier.
TABLE 21
REMOVAL OF TOTAL PHOSPHORUS (as P) DURING
EVALUATION OF THE LOW LIME PROCESS
mg/1 % Removal
Raw 6.63
Primary 19.31
Secondary 2.75 58.5
Lime Clarified .56 91.6
Filtered .29 95.6
Carbon Adsorption .16 97.6
Note:
Secondary Plant recycle enters between the raw sample
point and the primary clarifier.
55
-------
TABLE 22
REMOVAL OF NITROGEN COMPOUNDS* DURING
EVALUATION OF THE LOW LIME PROCESS
NH-
TKN
NO,
NO,
TOTAL N
Raw
Primary
Secondary
Lime Clarified
Filtered
Carbon Adsorption
7.92
9.64
9.11
7.60
8.87
10.37
9.92
8.34
5.02
3.67
4.38
5.16
.69
.64
.58
.48
14.59
14.68
14.88
13.98
All values mg/1 as N
56
-------
TABLE 23
LOADING RATES DURING EVALUATION OF THE LOW LIME PROCESS
Secondary Plant
Flow 4.858 mgd
Primary Clarifiers 2-60 ft dia. 859 gpd/sq ft
Secondary Clarifiers 2-70 ft dia. 631 gpd/sq ft
Model AWT Plant*
External Flow and Recycle = Total Flow
4.092 mgd + 0.588 mgd = 4.680 mgd
First Stage Reactor Clarifier 1-80 ft dia. 931 gpd/sq ft
Second Stage Reactor Clarifier 1-70 ft dia. 1216 gpd/sq ft
Dual Media Filtration (5 units)
242 sq ft/unit 3867 gpd/sq ft
or
2.69 gpm/sq ft
Carbon Adsorption
Column Set #1 Avg. Flow 1.285 mgd 5.05 gpm/sq ft
Column Set #2 Avg. Flow 1.117 mgd 4.39 gpm/sq ft
Column Set #3 Avg. Flow 1.095 mgd 4.31 gpm/sq ft
*
Based on flow plus internal recycle streams.
57
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TABLE 24
PLANT RECYCLE FLOWS DURING LOW LIME EVALUATION
Sources Total gal/day % of Flow
1. Filter Backwash
38,720 gal x 6 filters/day 232,320 4.96
2. Carbon Column Backwash
24,017 gal x 4 columns/day 96,068 2.05
3. Recalcination Furnace
164 gal/min x 1440 min/day 236,160 5.05
4. Misc. (centrate, pump sealing water,
flushing and wash water) 23,452 0.50
Total 588,000 12.56
All percentages are based on an average flow of 4.680 mgd.
TABLE 25
CHEMICAL USAGE IN THE LOW LIME PROCESS
Lime
Total Virgin Pounds 390,829
Average Available Lime Index (ALI) 87%
Available CaO, Ib 340,018
Average Daily Usage, Ib 4,097
Average dose, mg/1 105.3
Total Recalcined Pounds 37,088
Average Available Lime Index (ALI) 50.8%
Available CaO,lb 18,841
Average Daily Usage, Ib 224
Average Dose, mg/1 8-l
Ferric Chloride (EeCl3)
Total Pounds added to clarifier 68,646
Average Daily Usage, Ib 817
Average Daily Dose, mg/1 25.2
Polymer Usage (Centrifuge Only)
Total pounds 1,828
Average Daily Usage, Ib 26.1
Pounds Polymer/Ton of Dry Sludge 6.8
58
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during the test period was virgin pebble lime with an available lime index
of 87%. The precipitated lime was recalcined and recycled back into the
system for only a short time. The recalcined lime had an average available
lime index of only 50.8%. In the single stage low lime system, carbon di-
oxide is used to reduce the pH below the solubility point of calcium carbon-
ate, thus preventing a lime scale problem in the filters. The pH reduction
was. accomplished in three basins and allowed the process water to enter the
mixed media filters at pH 7.4. Again with the extra capacity built into
the system the reduction of pH from 10.5 to below 8.0 was easily controllable.
The carbon dioxide used during most of the project was purchased and was not
that produced by the recalcination furnace. The carbon dioxide used was 98%
pure CO2 which was diluted by the flue gas compressor to approximately 10%
CC>2 concentration.
The results from the chemical clarification system are presented in
Tables 19 through 22 and in Table 26. As shown in Table 21, the lime
clarification system reduced the phosphorus concentration to 0.56 mg/1
as P. This is an 80% capture of phosphorus that entered the clarifier
gystem.
TABLE 26
REMOVALS OF CHEMICAL OXYGEN DEMAND (COD)
AND TOTAL ORGANIC CARBON (TOC) DURING
EVALUATION OF THE LOW LIME PROCESS
COD
mg/1
Raw
Primary
Secondary
Lime Clarified
Filtered
Carbon Adsorption
32.54
25.45
17.03
9.34
TOC
mg/1 as C
16.0
11.6
9.6
3.6
Following the chemical clarification system, the effluent was fed to
dual media filters. Five filters were in continuous operation with one
additional unit either in backwash or standby. The flow as shown in Table 23
averaged 2.69 gpm/sq ft. The filters were backwashed on a 24-hour cycle
at an average headless of 8 feet.
59
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As seen in Table 20 the dual media filters removed suspended solids for
and average capture of 5.6 mg/1 per unit. As mentioned in the high lime
process discussion, an accurate efficiency of the filters could not be deter-
mined because of the undetermined portion of the calcium carbonate that was
solubilized during pH adjustment.
Following filtration the filtrate was adjusted to pH 7.4 by the addition
of carbon dioxide in the stabilization basin and pumped to the three parallel
activated carbon adsorption systems. As seen in Table 23, the hydraulic load-
ings of the three systems averaged 5.05, 4.39 and 4.31 gpm/sq ft, respective-
ly. The flows through the individual units were controlled successfully as
instrumentation was refined and additional equipment was added. The lead
column of each system was backwashed every 24 hours with the final columns
being backwashed every 48 hours. A total of 24,000 gallons of water was
required to backwash a carbon column as compared to 39,000 gallons to backwash
a dual media filter. During the backwash of a column, treatment by only one
column of that particular train was provided.
The performances of the activated carbon system based on the composite
samples of the three trains are presented in Tables 19 through 22 and in
Table 26. The results are the averages for the 89 days of operation. From
data in Table 26, the units removed approximately 63% of the inlet Total
Organic Carbon and 44% of the Chemical Oxygen Demand. The degree of nitri-
fication in the activated carbon columns, from 4.38 to 5.16 mg/1 of NO3~N,
(Table 22) was typical of operation of the units throughout the study and
was considerably lower than the nitrification produced in the high lime mode.
The performances of the individual carbon columns are presented in
Tables 27, 28 and 29. The loadings to the columns, based on the cumulative
flows through each column, are presented in Table 30. These loadings and
cumulative flows, except for column T-14, included the high lime process
evaluation. Column T-14 was regenerated during the plant shut down between
the high and low lime modes.
The waste solids from the clarifier were pumped to the gravity thicken-
er with a solids concentration generally between 1-5%. The concentration
varied as a function of flow which averaged 54,000 gal/day. As discussed in
the high lime process section, it was difficult to accurately sample and
analyze the sludge because of errors associated with the highly concentrated
waste stream. Therefore, the chemical analyses of the liquid streams were
considered to be more reliable because of the accuracy of the chemical analy-
ses in combination with good flow measurements. As shown in Table 31, the
average solids captured in the clarifier was 5509 Ib/day. The inerts in the
calculation include inert solids from the backwash cycles, centrate, re-
calcined lime and furnace operation.
60
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TABLE 27
PERFORMANCE OF CARBON ADSORBER TRAIN #1
DURING EVALUATION OF THE LOW LIME PROCESS
Average Flow 1.285 mgd
Cumulative flow to July 8, 1974
Total Organic Carbon (TOC)
Influent
Intermediate
Effluent
T-15
T-14
mg/1 as C
10.2
5.2
3.7
322.860 mil gal
122.953 mil gal
% Removed
49.0
63.7
Biochemical Oxygen Demand (BOD)
Influent
Intermediate
Effluent
mg/1
6.8
2.6
2.4
61.8
64.7
Chemical Oxygen Demand (COD)
Influent
Intermediate
Effluent
*
Note:
mg/1
18.1
10.9
7.8
39.8
56.9
Adsorber column T-14 was regenerated during January 1974 and
the sequence was switched to place adsorber T-15 in the lead.
Column T-14 operates as the number 2 unit.
61
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TABLE 28
PERFORMANCE OF CARBON ADSORBER TRAIN #2
DURING EVALUATION OF THE LOW LIME PROCESS
Average Flow 1.117 mgd
Cumulative flow to July 8, 1974*
Total Organic Carbon (TOC)
Influent
Intermediate
Effluent
T-16
T-17
mg/1 as C
10.0
5.4
4.2
244.329 mil gal
284.679 mil gal
% Removed
46.0
58.0
Biochemical Oxygen Demand (BOD)
Influent
Intermediate
Effluent
mg/1
6.6
2.5
2.2
62.1
66.7
Chemical Oxygen Demand (COD)
Influent
Intermediate
Effluent
mg/1
18.0
10.3
10.2
42.8
43.3
* Note:
Carbon adsorber T-16 was taken out of service for a short period
of time to repair a faulty valve.
62
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TABLE 29
PERFORMANCE OF CARBON ADSORBER TRAIN #3
DURING EVALUATION OF THE LOW LIME PROCESS
Average Flow 1.095 mgd
Cumulative flow to July 8, 1974
T-18 225.981 mil gal
T-19 225.981 mil gal
Total Organic Carbon (TOC)
Influent
Intermediate
Effluent
mg/1 as C
10.1
5.3
3.9
% Removed
47.5
61.4
Biochemical Oxygen Demand (BOD)
Influent
Intermediate
Effluent
mg/1
6.6
2.4
2.2
63.6
66.7
Chemical Oxygen Demand (COD)
Influent
Intermediate
Effluent
mg/1
18.3
12.4
8.8
32.2
51.9
63
-------
Set 1
T-15
T-14**
Set 2
T_16***
T-17
Set 3
T-18
T-19
TABLE 30
CUMULATIVE COD AND TOC LOADINGS
AT THE END OF THE LOW LIME EVALUATION*
Ibs COD/lb Carbon Cumulative Flow, mil gal Lb TOC/lb Carbon
.19299
.08128
.22266
.12755
.23164
.09940
322.860
122.953
244.329
284.679
225.981
225.981
.08594
.02605
.11206
.05709
.12044
.02576
* Loadings cumulative since columns were placed in service and include
the high lime evaluation except for T-14.
** Regenerated December 1973
*** T-16 was out of service due to a malfunctioning valve for approximately
one and one-half months.
TABLE 31
DAILY SOLIDS PRODUCTION FOR THE LOW LIME EVALUATION
Total solids production in the clarifier based on an average flow of 4.68 mgd
CaCO
Suspended Solids
Mg (OH) 2
Fe (OH)
Inerts
831 Ibs/mil gal
98.58 Ib/mil gal
22.52 Ib/mil gal
16.93 Ib/mil gal
116.6 Ib/mil gal
91.74 Ib/mil gal
X
X
X
X
X
X
4.68
4.68
4.68
4.68
4.68
4.68
TOTAL
= 3889 Ib/day
461 Ib/day
105 Ib/day
79 Ib/day
546 Ib/day
429 Ib/day
5509 Ib/day
64
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VIII. CARBON REGENERATION
With the failure of the two reactor clarifiers on November 18, 1973,
parbon column T-14 was prepared for regeneration. As seen in Tables 13
and 16, the carbon in column T-14 was not exhausted, but because the grant
period was ending it was felt that at least one regeneration should be done.
The column to date had averaged 57% removal of the Total Organic Carbon (TOC)
initially 7.5 mg/1 as C, to produce an effluent with 3.2 mg/1 as C. With
the main portion of the plant out of service, the regeneration system was
able to be started with full operator attention. This proved to be quite
necessary as numerous problems were encountered with the carbon transfer
system. Carbon adsorbers T-16 and T-18 were regenerated in August 1975,
and these data are also presented.
As seen in Table 32, one of the difficulties encountered in the regener-
ation process was obtaining an accurate measurement of the quantities of
carbon involved. The amount of carbon delivered by the supplier on the
original delivery for column T-14 was 72,100 pounds, however, based on
physical measurements and assuming a bulk density of 26 Ib/cu ft, it was
calculated that 73,000 pounds was delivered, a difference of 1.2%. Table 33
shows the operational data. Variations in the regeneration procedures were
numerous as means of transferring and measuring the carbon were being de-
veloped. The daily amounts of carbon in the spent and regenerated carbon
storage tanks were difficult to determine because of uneven carbon levels
and the inaccessibility of the tanks. An accurate measurement of the losses
due to regeneration could therefore not be determined with the total system
and measuring options available. The closest estimate of losses that can
be given from the three regenerations is that losses were between 8 and 10%.
TABLE 32
INVENTORY OF CARBON
T-14 T-16 T-18
Carbon reported delivered by supplier, Ib 72,100 70,640 69,640
Carbon as measured in column,Ib* 73,000 69,430 68,888
Carbon before regeneration, Ib* 69,792 69,800 68,500
Carbon fed to the furnace based on feed rate,lb 65,112 67,200 61,500
*Note, calculations based on 26 Ib carbon/cu ft.
65
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As compared in Tables 33 through 35, the regenerations of columns T-16
and T-18 resulted in lower iodine numbers in the regenerated carbon because
of higher loadings for the spent carbon and increased feed rates to the re-
generation furnace. The regeneration process was controlled by the results
from the apparent bulk density test. This test was the only one in which
the results could be received immediately and corrections made to the process.
As can be seen in Table 34, the densities also reflected the increased feed
rates and resulting decreased regeneration efficiency. Iodine tests along
with determination of ash concentrations were performed in the laboratory
and normally required 5-7 days before the results were available. These
numbers, therefore, could not be used for timely adjustment of operating
parameters for carbon regeneration. Table 36 shows one sieve analysis from
column T-18 indicating some breakdown in particle size due to the regener-
ation and transfer processes. Table 37 gives the natural gas consumption
in the three regenerations and the maximum temperatures through which the
carbon passed during regeneration. The hottest hearth was the bottom hearth,
#4.
The three carbon columns which were regenerated represent the startup
and debugging phases of operation. Numerous modifications and corrections
would have to be made to the delivery and measuring operations before carbon
losses could be accurately determined.
TABLE 33
OPERATING DATA FOR REGENERATION OF CARBON
T-14
T-16
T-18
Total feed to furnace, Ib
Total operating days
Average hours of feed/day
Carbon feed rate, Ib/min
Transfer rate, Ib carbon/gal water
Steam rate, Ib/hr
Ib steam/lb carbon
65,112 67,200 61,500
16 15 12
16.3 9.3 10.8
4.14
0.47
140
.56
8.03
.51
146
.30
7.9
1.0
135
.28
66
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TABLE 34
RESULTS OF LABORATORY ANALYSES OF CARBON
Iodine number, virgin
Iodin,q number, spent
Iodine number, regenerated
Ash number, spent, %
Ash number, regenerated, %
Apparent bulk density, virgin, grams/cc
spent, grams/cc
regenerated, grams/cc
TABLE 35
TOG loading (X10
BOD loading (X10
COD loading (X10
-2 Ib TOC j
Ib carbon
-2 Ib BOD )
Ib carbon
-2 Ib COD )
Ib carbon
T-14
950
748
936
5.0
6.1
.500
.552
.486
2GENERA'
T-14
9.6
2.4
21.1
T-16
950
428
741
8.2
8.9
.500
.572
.523
riON
T-16
18.5
16.5
41.0
T-18
950
554
747
7.3
9.3
.500
.565
.516
T-18
19.5
18.9
46.3
67
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TABLE 36
SIEVE ANALYSES OF CARBON FROM T-18 CARBON ADSORBER
Sieve size /U.S. No.
4
8
10
12
14
16
20
30
40
325
Spent (% retained)
2.60
9.99
17.71
19.49
19.67
12.96
11.96
4.37
.70
.55
Regenerated (% retained)
0
4.25
24.69
21.77
17.71
11.28
11.08
3.07
.40
.75
TABLE 37
FURNACE CONDITIONS DURING CARBON REGENERATION
T-14 T-16 T-18
Natural gas actually used, cu ft/hr 1494 1465 1512
Temperature (max.hearth)(°F) 1677 1718 1800
Temperature (afterburner) (°F) 1375 1344 1215
Ratio of natural gas to carbon 6.0 3.0 3.2
cu ft natural gas/lb of carbon
68
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IX. COST ANALYSIS
The major purpose of this project as conceived by EPA and WSSC was to
gather reliable data with regard to operating costs and operational problems.
Most of EPA's previous work in the Physical-Chemical AWT processes had been
accomplished on a small pilot scale and the purpose of the Piscataway Model
Plant project was to build and evaluate a larger system. The following pages
show cost breakdowns and mention several operating problems encountered.
Costs for the Model Plant were high for the following reasons:
1. Data are based on the startup period where equipment was
being modified and adjusted.
2. Operators were unfamiliar with the plant. The majority
of the staff had little training in the wastewater field
or prior experience in treating wastewater.
3. Processes were not operated within optimum ranges of
efficiency.
The caliber and quality of the operating staff should be a major con-
sideration when discussing and selecting unit processes for the facility.
The startup period at the Piscataway Model AWT Plant included numerous prob-
lems revolving around three unfortunate situations. First was the inability
to hire and retain qualified personnel. Secondly, numerous mechanical prob-
lems were further complicated when a fire broke out late in the construction
phase, causing extensive smoke damage to most of the pneumatic controls and
automatic equipment, as well as creating delays in completion of construction.
Third was the lack of redundancy in key pieces of equipment.
The basis for design was that the plant would be a model facility and
have a five-year design life. This was done with the concept of obtaining
as much technical data as possible while keeping the capital cost at a
minimum. This, however, resulted in design deficiencies that caused oper-
ating problems that might not occur in a plant designed for a longer life
and for greater operating efficiency. Prior to startup in late 1972, the
new engineer and his staff of ten plant operators attended daily classroom
training sessions and performed actual plant work before entering the Model
Plant. This period included spending more than 40 hours at the EPA Blue
Plains Pilot Plant to gain vital operating experience. The operators were
also given practical experience while working at the Piscataway Secondary
Treatment Plant. Although these eleven persons began the initial testing
and startup in September 1972, the plant had lost the trained engineer and
7 of the original 10 operators by December 1972. During the next year a
total of 51 operators and 3 engineers-in-charge had been assigned to the
plant at one time or another. This high turnover combined with both normal
construction and design defects resulted in a long startup period. One of
the major lessons learned from the Model Plant testing was that quality
69
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personnel are needed if a relatively complex plant is to be operated on a
continuous basis.
Capital Expenditures
As mentioned previously, the concept of a Model Plant began with EPA
in 1966. The original talks were based on what size plant could be built
for approximately 2.5 million dollars. However, following a preliminary
design study it was determined that a 5 mgd plant could be constructed for
3.2 million dollars and the project proceeded into design and construction.
As shown in Table 38, the final cost of the Model Plant was 4.68 million
dollars. Table 39 shows a more detailed breakdown of the costs in which the
EPA's final share was 3.1 million dollars, the State of Maryland paid 0.3
million and WSSC paid 1.1 million dollars.
Tables 40 and 41 give a cost breakdown for the Piscataway Secondary
Plant. Phase I was for the 5 mgd system which is used to treat the sewage
entering the Model Plant. The total cost of the 30 mgd Secondary Plant was
13.5 million dollars. Table 42 gives the breakdown on engineering services
supplied to EPA and WSSC by Roy F. Weston, Inc.
Tables 43 through 45 give breakdowns of the Model Plant expenses for
the unit processes. Note that in Table 43 an effort has been made to esti-
mate the costs of a 5 mgd plant if the low lime concept were to be the basis
for design. Also in Table 43, note that the engineering services have been
included as complete plant costs and are not broken down for the individual
processes.
TABLE 38
CAPITAL COSTS OF THE MODEL PLANT
General Contractor (Main Contract) $3,037,100
Change Orders 210,338
Furnace Contractor 529,000
Activated Carbon 148,356
Centrifuges 66,480
Electric Substation 153,800
Sub Total $4,145,074
Engineering Services 535,243
Total $4,680,317
70
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TABLE 39
DISTRIBUTION OF CAPITAL COSTS OF THE MODEL PLANT
U.S. Environmental Protection Agency
Under Project 17080DZY
Research and Development EPA's share 75%
$3,200,000 x .75 = $2,400,000
Under Project WPC - MD - 233
Additional Facilities $1,320.453
55% EPA's share
$1,320,453 x .55 = $ 726,249
25% State of Md. share
$1,320,453 x .25 = $ 330,113
Washington Suburban Sanitary Commission
Main Contract $ 800,000
Additional Facilities $ 264,090
$1,064,090
71
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TABLE 40
WSSC CONSTRUCTION PHASES FOR THE PISCATAWAY SECONDARY TREATMENT PLANT
-j
N)
PHASE
I
II
III
IV
V
COST
$
$
$
$
$
3
2
2
1
1
,027
,510
,431
,079
,702
,000
,435
,634
,338
,500
CONSTRUCTION
BEGUN
1966
1968
1970
1970
1972
CONSTRUCTION
COMPLETED
1967
1970
1974
1971
1974
REMARKS
5 mgd complete secondary plant and
handling facilities.
25 mgd expansion of aeration basin
secondary clarifiers.
solids
and
25 mgd expansion of grit chamber, primary
clarifier and thickening areas.
Four polishing ponds installed improved
process effluent.
Construction of additional solids h
tand ling
VI $ 1,768,762
VII
$ 967,317
1972
1973
1974
1975
facilities including a vacuum filter and
three incinerators.
Construction of additional solids handling
including vacuum filter and incinerator.
(In conjunction with Phase III).
Modification to existing plant including
samplers and misc. safety equipment.
TOTAL $13,486,986
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TABLE 41
COST BREAKDOWN FOR THE
30 MGD SECONDARY TREATMENT PLANT
Project I
5 mgd
Primary
Secondary
Solids Handling
Project II
Ponds
Other
Project III
Expansion 30 mgd
WSSC Phases
Primary
Secondary
Solids Handling
I
II
III
IV
V
VI
VII
$ 774,078
$ 1,501,907
$ 751,015
Total Cost
775,621
302,518
Total Cost:
$ 1,495,313
$ 2,748,209
$ 4,796,426
Total Cost
$
$
$
$
$
$
$
3,027,000
2,510,435
2,431,634
1,079,338
1,702,500
1,768,762
967,317
$ 13,486,986
$ 3,027,000
$ 1,079,339
$ 9,039,948
$ 13,146,287
Note: Difference between two totals is $340,669 which includes expenses
for the Model Plant, miscellaneous landscaping, and road work.
73
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TABLE 42
COSTS OF ENGINEERING SERVICES FOR THE MODEL PLANT
PRELIMINARY ENGINEERING DESIGN AND REPORT $ 28,449
STUDY OF ADVANCED TREATMENT 7,281
PREPARATION OF PLANS AND SPECIFICATIONS 288,511
CONSTRUCTION SERVICES 190,095
PREPARATION OF OPERATIONS MANUAL 20,907
Total $535,243
TABLE 43
CAPITAL COSTS FOR THE MODEL PLANT UNIT PROCESSES
Process High Lime Cost Low Lime Cost*
Lime Clarification $1,340,190 $ 991,000
Filtration & Carbon Adsorption 1,150,354 1,150,354
Solids Handling 1,099,247 291,930
Carbon Regeneration 401,483 401,483
Sub Total $3,991,274 $2,834,767
Electrical Substation 153,800 ** 109,198
Engineering Services 535,243 ** 380,022
Total $4,680,317 $3,323,987
*Low lime cost is based on a calculated estimate.
**Both of these numbers are proportional to the sub totals.
74
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TABLE 44
BREAKDOWN ON CAPITAL COSTS FOR THE MODEL PLANT UNIT PROCESSES
Clarification
Excavation
Concrete Work
Mechanical Work
Electrical Work
Painting
Other Work
Filtration
Steel Shelves
Underdrains
Filter Media
Mechanical Work
Piping and Valves
Painting
Electrical Work
Other
Carbpn Adsorbers
Steel Tanks and Supports
Filter Bottom
Mechanical Work
Piping and Valves
Electrical Work
Painting
Other Work
Operations Building
Centrifuge Area & Cake Handling System
(excluding centrifuge)
Lime Handling System
FeCl_ and Polymer System
•J
Motor Contro^ Center and Power Distribution
Ins trumenta tion
Miscellaneous Equipment, Steel and Clean-up
52,700
187,400
237,000
28,000
10,000
20,300
535,400
118,000
35,000
8,000
29,500
32,000
3,000
600
7,100
232,200
$ 232,000
28,000
26,000
75,000
6,000
7,900
1,800
$ 367,700
$ 1,076,900
33,800
100,200
43,200
105,500
308,800
226,337
$ 3,031,037
75
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TABLE 45
CAPITAL COST BREAKDOWN FOR THE MODEL PLANT EQUIPMENT
GENERAL CONTRACT
Reactor Clarifier
Recarbonation Basin & Associated Tanks
Lime Pumping Stations
Thickener & Pumping Station
Operations Building
Dual Media Filters
Backwash Tanks
Stabilization Tank
Activated Carbon Adsorbers (excluding carbon)
Carbon Storage Tanks
Lime Bins
Lime Slaker System
Assorted Lime Handling Systems
Lime Cake Handling Systems
Miscellaneous Plant Equipment
Chemical Feed System (Ferric & Polymer)
Steam Generator
Platform & Structures
Motor Control Center & Power Distribution
Instrumentation
Sitework & Piping
Sub Total
Change Orders
Furnace Contracts
Activated Carbon
Centrifuges
COST
$ 394,000
92,400
81,000
68,000
1,076,900
194,000
78,400
50,500
247,500
30,500
31,500
25,700
43,000
33,800
61,600
43,200
13,500
69,200
105,500
308,800
88,100
$ 3,037,100
210,338
529,000
148,356
66,480
DATE OF
COMPLETION
October 72
October 72
October 72
October 72
April 72
April 72
April 72
April 72
April 72
April 72
May 72
February 73
February 73
October 72
April 72
October 72
April 72
April 72
March 72
February 73
April 72
February 73
March 73
i
October 72
July 72
Total
$ 3,991,274
76
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Operational Costs
The following pages provide operating costs based on pay schedules
shown in Table 46 and cost of chemicals and energy shown in Table 47.
Several factors to consider when analyzing these data are:
1. The miscellaneous group used in the tables and distribution diagrams
includes items such as water usage, hoses, shovels, office and lab
supplies, and cleaning supplies.
2. All chemical costs are FOB at the plant.
3. Costs do not include those for laboratory personnel and for the
analytical tests.
4. No overhead costs are included in the cost breakdowns other than
those for the staff on hand.
As shown in Table 48, operating costs for tertiary treatment using the
Low Lime Process with wasting of wet solids were 28.92C/1000 gallons ($289/
mil gal). Table 49 shows an operating cost of 32.97C/1000 gallons for the
Low Lime Process with solids dried and wasted and Table 50 shows an operat-
ing cost of 35.80C/1000 gallons for the High Lime Process. The costs reflect
the plant operation, but do not consider the cost to landfill the waste from
the three types of lime treatment. With the three types of processes, the
personnel requirements are the same. In certain instances, and depending
on unit processes, personnel requirements can be cut for the Low Lime Proc-
esses. At the Piscataway Model Plant, the staff operated separately from
the staff at the Piscataway Secondary Plant. This separation of the staffs
caused increased costs. Table 51 gives the personnel breakdown in the
various unit processes.
Figures 15 through 17 are breakdowns of the operating costs as present-
ed in Tables 48 through 50. Figures 18 through 21 give distribution dia-
grams of the capital costs and power costs of the Model Plant. As mentioned
previously, the calculations for the Low Lime Process in both the capital
cost and power requirements are the best estimates that can be made from
the available data.
77
-------
LIME
CLARIFICATION
44.7%
SOLIDS
HANDLING
16.6%
CARBON
REGENERATION
5.8%
FILTRATION
8.7%
REPAIR
PARTS
20.8%
MISC.
3.4%
Figure 15. Distribution of operating costs for the low
lime process with wasting of wet solids.
78
-------
LIME
CLARIFICATION
31%
SOLIDS
HANDLING
29%
CARBON
REGENERATION
5.0%
ILTRATION
AND
CARBON
ADSORPTION
7.6%
MISC. 3.0%
REPAIR
PARTS
24.4%
Figure 16. Distribution of operating costs for the low
lime process with solids dried and wasted.
79
-------
LIME
CLARIFICATION
31%
SOLIDS
HANDLING
32%
CARBON
REGENERATION
4.7%
REPAIR
PARTS
22.5%
FILTRATION
AND
CARBON
ADSORPTION
7%
MISC.
2.8%
Figure 17. Distribution of operating costs for the high
lime process.
80
-------
LIME
CLARIFICATION
29.8%
FILTRATION
AND
CARBON
ADSORPTION
34.6%
ENGINEERING
11.4%
SOLIDS
HANDLING
8.8%
CARBON
REGENERATION
12.1%
Figure 18. Distribution of capital costs for the low
lime process.
81
-------
LIME
CLARIFICATION
28.6%
FILTRATION
AND
CARBON
ADSORPTION
24.6%
ENGINEERING
11.6%
SOLIDS
HANDLING
23.5%
Figure 19. Distribution of capital costs for the high
lime process.
82
-------
LIME
CLARIFICATION
68.0%
FILTER
AND
CARBON
ADSORPTION
23.4%
SOLIDS
HANDLING
2.6%
MISC. 4.9%
Figure 20. Distribution of power requirements for the
low lime process.
83
-------
LIME
CLARIFICATION
58.5%
SOLIDS
HANDLING
20.9%
FILTER
AND
CARBON
ADSORPTION
16.3%
MISC. 3.4%
Figure 21. Distribution of power requirements for the
high lime process.
84
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TABLE 46
PAY SCHEDULES AT PISCATAWAY
Effective July 1, 1974
Yearly Pay
Engineer II $14,211 - $20,956
Engineer I $12,316 - $17,317
Senior Wastewater Plant Operator $11,710 - $14,946
Wastewater Plant Operator $ 9,050 - $14,230
Laborer/Driver $ 8,507 - $10,630
Chemist II $15,410 - $19,195
Chemist I $11,710 - $16,491
Laboratory Technician $ 8,561 - $14,230
Bacteriologist I , $11,710 - $16,491
Electrical/Mechanical Engineers $12,920 - $24,691
Asst. Electrical Maintenance Supervisor $15,419 - $19,195
Electrical Mechanic $10,379 - $13,853
Electrical Mechanic Apprentice $ 9,547 - $11,560
TABLE 47
COST FIGURES FOR ENERGY AND CHEMICALS
Chemical
Pebble Lime $ 22.15/ton
Ferric Chloride $128.50/ton
Polymer $ 1.35/lb
Carbon Dioxide $ 60.00/ton
Power
Electricity $0.0185/kWh
#2 Fuel Oil $0.2545/gal
Natural Gas $2.50/first 5 therms*
$0.175/second 10 therms
$0.165/next 15 therms
$0.148/next 500 therms
Miscellaneous
Water $0.48/1000 gal
Granular Carbon $826.61/100 cu ft
CO2 Tank Rental $265/month
* 1 therm = 100 cu ft
85
-------
TABLE 48
Operating Costs for the
Low Lime Process With
Wasting of Wet Solids
Lime Clarification
Solids Handling
Filtration & Carbon Adsorption
Miscellaneous (Water, supplies)
Repair Parts
Carbon Regeneration
TABLE 49
Operating Costs for the
Low Lime Process With
Solids Dried & Wasted
Lime Clarification
Solids Handling
Filtration & Carbon Adsorption
Miscellaneous (Water, supplies)
Repair Parts
Carbon Regeneration
Cost/1000 Gallons
12.92*
4.80*
2.52*
1.00*
6.00*
1.68*
28.92*/1000 gallons
Cost/1000 Gallons
10.22*
9.55*
2.52*
1.00*
8.00*
1.68*
32.97*/1000 gallons
86
-------
TABLE 50
Operating Costs for the
High Lime Process
Lime, Clarification
Solids Handling
Filtration s Carbon Adsorption
Miscellaneous (Water, supplies)
Repair Parts
Carbon Regeneration
Cost/1000 Gallons
11.08$
11.52C
2.52$
l.OOC
8.00C
1.68C
35.80^/1000 gallons
TABLE 51
Personnel Breakdown
by Unit Processes
Lime Clarification
Solids Handling
Filtration & Carbon Adsorption
Carbon Regeneration
Total
1 Engineer
15 P3,ant Operators
2 Laborers
18 Total Staff
Operators
hr/day
27
34
6
12
Engineer
hr/day
2
3
1
2
Laborers
hr/day
6
8
0
2
79
16
87
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APPENDIX
CALCULATION OF RECALCINATION FURNACE HEAT BALANCE
Feed rate of wet sludge
Moisture content by weight
Total solids by weight
CaCO concentration in dry solids
% volatile in dry solids
Mg (OH) concentration in dry solids
OH (P04)
concentration in dry solids
Given: Average temperature based on 41 days operation
1. Exhaust gas temperature
2. Calcined product temperature
3. Furnace shell temperature
4. Feed inlet temperature
5. Center shaft cooling air temperature
Heat Balance Calculations
3315 Ib/hr
61.5%
38.5%
72.0%
8.0%
12.0%
9.0%
910°F
1000°F
175°F
60°F
235°F
Equations
Ib/hr
Ib/hr
Ib/hr
*Ib/hr
*Ib/hr
Ib/hr
**lb/hr
* Note: Chemical Equation
HO 3315 Ib/hr x 61.5% = 2038.7
£
Solids 3315 Ib/hr x 38.5% = 1276.3
CaC03 1276.3 Ib/hr x 72% = 918.9
CaO 918.9 Ib/hr x 56/100= 514.6
CO 918.9 Ib/hr x 44/100= 404.3
£t
Volatile 1276.3 Ib/hr x 8% = 102.1
Inert Ash 1276.3 Ib/hr x 20% = 255.3
CaCO,
CaO + CO,
3 2
100 =56+44 molecular weight
** Note: Inert Ash is made up of Mg, Fe, and P Compounds.
88
-------
Summary of Feed Composition
Constants
1. m = mass quantity in Ib/hr
2. cp = heat capacity (specific heat)
3. H = -latent heat of vaporization
I. Heat Demand - material plus reaction requiring heat input
A. Heat required to heat up HO, evaporate HO, superheat water vapor
to 910°F
1. Heat up liquid water to 212 F
H20 liquid 60°F -> H20 liquid 212°F
Q = m C AT
p
= 2038.7 Ib/hr x 1.0 BTU x (212°F - 60°F)
lb°F
= 309,882 BTU/hr
2. Evaporate water at 212 F
HO liquid -»• HO steam
Q = m H
= 2038.7 Ib/hr x 970 BTU
Ib of HO
= 1,977,539 BTU
hr
3. Superheat steam to exhaust gas temperature
H20 steam 212°F -»• HO steam 1000°F
Q = M c AT
= 2038.7 Ib/hr x .5 BTU x (1000°F - 212°F)
lb°F
= 803,248 BTU/hr
Total heat requirement of water
309,882
1,997,539
803,248
3,090,669 BTU
hr
B. Heat required to break CaC03 solid -> CaO solid + C°2
Q = m H
= 918.9 Ib/hr x 785 BTU
Ib CaCO
= 721,336 BTU
hr
89
-------
C. Heat required to heat up product CO to 910 F
Q = m- c AT
2 p
= 404.3 Ib/hr x .25 BTU/lb°F x (910° - 60°F)
= 85,914 BTU
hr
D. Heat required to heat up CaO and ash to calciner outlet temperature
0 = rri c AT
3 p
= (514.6 + 255.3) Ib/hr x .3 BTU x (1000°F - 60°F)
Ib^F
= 769.9 Ib/hr x .3 BTU
Ib75!1
= 217,112 BTU/hr
E. Radiation loss
1. Furnace Shell Diameter 18.75 feet
Height 22.00 feet
Total surface area of furnace = 1852 sq ft
Average heat loss at 175 F skin temperature is 250 BTU
hr sq ft
Q = 250 BTU x 1852 sq ft
hr sq ft
= 463,000 BTU
hr
2. From center shaft cooling air
60°F - Mr 235°F
Q = m * c AT
4 p
= 3500 cu ft/min. x 60 min/hr x .02 BTU x (235°F - 60°F)
= 735,000 BTU CU ft
hr
* Note: air volume through center shaft
Total radiation loss - 463,000 + 735,000 = 1,198,000 BTU
hr
Total required heat input
Heat (H20) 3,090,669
Endothermic 721,336
Product C02 85,914
Lime and Ash 217,112
Radiation 1,198,000
5,313,031 BTU
hr
90
-------
II. Heat Release
1. Volatile release 102.1 Ib/hr x 10,000 BTU = 1,021,000 BTU
Ib hr
2. Heat required from fuel oil
Total heat required 5,313,031 BTU
hr
Heat release volatile material 1,021,000 BTU
hr
4,292,031 BTU
hr
#2 fuel oil thermal output = 138,000 BTU/gal*
assuming 100% efficiency
* Reference: North American Combustion Handbook
Assuming 20% excess air and a flue gas outlet temperature of 910 F,
the percent of gross fuel input which is available is 68%.
Therefore, total gallons of #2 fuel oil required to produce
4,292,031 BTU is equal to: 4,292,031 BTU = 45.74 gallons
hr hr hour
138,000 BTU x 0.68
gal
45.74 gal x 24 hr/day = 1098 gal/day
hr
Actual fuel oil usage was 1169 gal/day
Therefore, the efficiency of operation is:
1098 x 100
1169
91
-------
TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
\. REPORT NO.
EPA-600/2-78-172
2.
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
OPERATIONAL RESULTS FOR THE PISCATAWAY
MODEL 5 MGD AWT PLANT
5. REPORT DATE
September 1978 ("Issuing Date^)
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
Thomas P. O'Farrell, Robert A. Menke
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Washington Suburban Sanitary Commission
Hyattsville, Maryland 20781
10. PROGRAM ELEMENT NO.
BC611
11. CONTRACT/GRANT NO.
S802943
12. SPONSORING AGENCY NAME AND ADDRESS
Municipal Environmental Research Laboratory-Cinti, OH
Office of Research and Development
U.S. Environmental Protection Agency
Cincinnati. Ohio 45268
13. TYPE OF REPORT AND PERIOD COVERED
IT) -
147SPbNSb'RiNG AGENCY tODE
EPA/600/14
15. SUPPLEMENTARY NOTES
Project Officer: Sidney A< Hannah,(513/684-7651)
16. ABSTRACT
A 5 mgd tertiary wastewater treatment plant was constructed to demonstrate
treatment of effluent from a 5 mgd step aeration activated sludge plant. The two-
stage high lime process with intermediate recarbonation, filtration and activated
carbon adsorption operated at the design rate for 36 days between two failures of
the reactor clarifiers. A single-stage low lime process with filtration and
activated carbon adsorption operated for 89 days. The combined secondary and
tertiary treatment removed > 97% of BOD, TSS and P in the raw wastewater. Capital
cost of the 5 mgd two-stage high lime system was 4.7 million dollars and operating
costs were estimated as 36 cents per 1000 gallons of wastewater.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Sewage Treatment*, Chemical Removal
(Sewage Treatment)*, Activated Carbon
Treatment, Coagulation, Clarification,
Filtration, Water Pollution
Physical-Chemical Treat-
ment, Tertiary Treatment
13B
18. DISTRIBUTION STATEMENT
RELEASE TO PUBLIC
19. SECURITY CLASS (ThisReport)
UNCLASSIFIED
21. NO. OF PAGES
106
20. SECURITY CLASS (Thispage)
UNCLASSIFIED
22. PRICE
EPA Form 2220-1 (Rev. 4-77)
92
* U.S. GOVERNMENT POINTING OFFICE: 1978— 6 57-060 /1477
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