WATER POLLUTION CONTROL RESEARCH SERIES
17050 DNW 02/72
ACTIVATED SLUDGE PROCESSING
U.S. ENVIRONMENTAL PROTECTION AGENCY
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WATER POLLUTION CONTROL RESEARCH SERIES
The Water Pollution Control Research Series describes
the results and progress in the control and abatement
of pollution in our Nation's waters. They provide a
central source of information on the research, develop-
ment, and demonstration activities in the water research
program of the Environmental Protection Agency, through
inhouse research and grants and contracts with Federal,
State, and local agencies, research institutions, and
industrial organizations.
Inquiries pertaining to Water Pollution Control Research
Reports should be directed to the Chief, Publications
Branch (Water), Research Information Division, R&M, En-
vironmental Protection Agency, Washington, D. C. 20460.
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CONTINUED EVALUATION OF OXYGEN USE IN
CONVENTIONAL ACTIVATED SLUDGE PROCESSING
by
Union Carbide Corporation
Linde Division
P.O. Box 44
Tonawanda, New York 14150
for the
OFFICE OF RESEARCH AND MONITORING
ENVIRONMENTAL PROTECTION AGENCY
Project # 17050 DNW
Contract # 14-12-867
February, 1972
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EPA REVIEW NOTICE
This report has been reviewed by the Environmental Protection
Agency and approved for publication. Approval does not signify
that the contents necessarily reflect the views and policies of
the Environmental Protection Agency, nor does mention of trade
name or commercial products constitute endorsement or recommenda-
tion for use.
11
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ABSTRACT
A process"- for treating municipal wastewater using high-purity oxygen
in the activated sludge process was further evaluated at Batavia, New
York in a full scale wastewater treatment plant. The process and
equipment used for treatment of the wastewater were similar to that
utilized in 1969 to demonstrate the feasibility of using high-purity
oxygen effectively and economically. This initial work has been reported
in Water Pollution Control Research Series Report 17050DNW05/70. The
present evaluation places emphasis on the production rate, filtration,
and digestion characteristics of oxygenated waste activated sludge.
Operation was conducted at two different system loadings over a period
of three months. Sludge production data obtained during each period
of operation was verified by performing solids and mass balances over
the entire system on a continuous basis. The mass balances indicate
that a valid accounting of solids profiles was obtained during the
operational periods. The excess solids production data obtained in this
contract generally confirm the observations made during the initial con-
tract that the Batavia oxygenation system produces 40-45% less biologi-
cal solids than the conventional Batavia air aeration system.
The filtering characteristics of waste activated sludge from an oxygen-
ation system were investigated. Waste sludge from the oxygenation
system clarifier underflow was filtered directly on a pilot-scale vacuum
filter. A filter cake was produced which disengaged easily, and had a
solids content of 15-24 percent. Filter yields as high as 4.5 pounds
per hour per square foot of filter area were obtained.
Waste activated sludge was aerobically digested in 800-gallon batches
using high-purity oxygen. The results of the batch studies indicate
that a 20-30 percent reduction in volatile suspended solids can be
obyained in 7-9 days. An essentially odorless stabilized sludge is
produced.
The economics of oxygenation systems as compared with conventional
diffused air aeration systems were revised from the cost estimates
presented in the initial contract work. The revisions reflect the
sludge production figures for both the initial and the present contract
work, as well as current cost estimates for an oxygenation system.
These economic studies indicate that costs with oxygenation are comparable
to costs with conventional diffused air aeration for a plant size of
1 MOD. As plant size increases, costs with oxygenation become less than
costs with conventional diffused air aeration with savings of about 20%
projected at this time for a 100 MGD design.
This report was submitted in fulfillment of Project Number 17050 DNW,
Contract 14-12-867, under the sponsorship of the Environmental
Protection Agency.
"Process currently marketed by Union Carbide Corporation under the name
UNOX.
111
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CONTENTS
Section Page
I CONCLUSIONS !
II RECOMMENDATIONS 3
III INTRODUCTION 5
Review of 1969 Oxygenation Contract Results 5
Discussion of Significant Technical Results from 1969
Oxygenation Contract 6
IV CONTRACT PROGRAM PLAN 17
V EQUIPMENT DESCRIPTION AND EVALUATION METHODS 19
Modifications to Oxygenation System ^
Data Acquisition and Analysis 20
Data Reduction 21
Solids and Mass Balance Methods 22
VI SOLIDS PROFILES AND EXCESS SLUDGE PRODUCTION 29
Operating Conditions - Phase I 29
Operating Conditions - Phase II 35
Mass Balances 39
Excess Sludge Production ^
Miscellaneous Results ^6
Economic Comparison of Oxygenation and Conventional
Diffused Air Aeration Treatment 52
VII VACUUM FILTRATION 61
Oxygenation System Waste Activated Sludge 64
Comparison of Oxygenated and Air Aerated Waste
Activated Sludges 71
Anaerobically Digested Sludge 74
Aerobically (Oxygen Aerated) Digested Waste
Activated Sludge 76
VIII AEROBIC DIGESTION 77
Solids Reduction 79
Oxygen Uptake Rates 85
Nutrient Profiles 85
IX ACKNOWLEDGEMENTS 89
X REFERENCES 91
XI APPENDICES 93
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FIGURES
No.
1 SOLIDS PRODUCTION CALCULATED FROM POON'S DATA 13
2 FLOW DIAGRAM FOR BATAVIA OXYGENATION SYSTEM MASS BALANCE ANAL. 25
3 BATAVIA TREATMENT SYSTEM PHASE I OPERATION 1970 CONTRACT 30
4 BATAVIA TREATMENT SYSTEM PHASE II OPERATION 1970 CONTRACT 36
5 COMPARATIVE CORRELATION OF EXCESS VSS AND MLVSS WITH
AERATION DETENTION TIME FROM INITIAL CONTRACT REPORT
17050DNW 05/70 ^
6 COMPARATIVE CORRELATION OF BOD REMOVED AND MLVSS WITH
AERATION DETENTION TIME FROM INITIAL CONTRACT REPORT
17050DNW 05/70 <+8
7 CORRELATION OF EXCESS VSS FROM AIR AERATION AND OXYGEN-
ATION SYSTEMS WITH BOD REMOVAL AND MLVSS FROM INITIAL
CONTRACT REPORT 17050DNW 05/70 49
8 EXCESS VSS FORMATION CORRELATIONS AND 95% CONFIDENCE
LIMITS FOR PREDICTED VALUES 1969 AND 1970 BATAVIA
OPERATIONS 50
9 TOTAL TREATMENT COSTS - 1 MGD TREATED 55
10 TOTAL TREATMENT COSTS - 6 MGD TREATED 56
11 TOTAL TREATMENT COSTS - 30 MGD TREATED 57
12 TOTAL TREATMENT COSTS 100 MGD TREATED 58
11 FLOW SHEET OF VACUUM FILTER APPARATUS 62
U OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE CAKE
MOISTURE VS. CYCLE TIME 66
15 OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE CAKE
YIELD VS. % SOLIDS IN FEED 69
16 OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE CAKE
YIELD VS. CYCLE TIME 70
17 OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE OBSERVED
CAKE YIELD VS.(t)-°-656 • (C.)°-767 72
18 OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE CALCULATED
VS.OBSERVED CAKE YIELD 73
vi
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FIGURES CONTD.
No.
19 OBSERVED CAKE YIELD FROM VACUUM FILTRATION OF MILWAUKEE'S
CONVENTIONAL THICKENED WASTE ACTIVATED SLUDGE VS.
CALCULATED CAKE YIELD OF OXYGENATED SLUDGE 75
20 REDUCTION OF READILY BIODEGRADABLE VOLATILE SUSPENDED
SOLIDS VS. DETENTION TIME 80
21 REDUCTION OF TOTAL VOLATILE SUSPENDED SOLIDS VS. DETENTION
TIME 82
22 REDUCTION OF TOTAL SUSPENDED SOLIDS VS. DETENTION TIME 83
23 RATIO OF VOLATILE SUSPENDED SOLIDS TO TOTAL SUSPENDED
SOLIDS VS. DETENTION TIME 84
24 OXYGEN UPTAKE RATE VS. VOLATILE SUSPENDED SOLIDS 86
vii
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TABLES
No. Page
1 COMPARATIVE SUMMARY OF AIR AERATION AND OXYGENATION
SYSTEM PERFORMANCE -1969 CONTRACT ?
2 SUMMARY OF DATA FROM UNION CARBIDE PILOT PLANT
PROGRAMS 15
3 AVERAGE OPERATING CONDITIONS PHASE I OPERATION
(8/30/70 10/25/70) 32
4 AVERAGE OPERATING CONDITIONS PHASE II OPERATION
(11/1/70 11/30/70) 37
5 SUMMARY OF SOLIDS AND MASS BALANCES FOR BATAVIA
OXYGENATION SYSTEM 41
6 AVERAGE WEEKLY BATAVIA MASS BALANCES FOR PHASES I
AND II 43
7 REVISED SUMMARY OF SLUDGE DISPOSAL ESTIMATES FOR AIR
AERATION AND OXYGENATION SYSTEM ECONOMIC COMPARISONS-
100 MGD PLANT 54
VIII
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SECTION I
CONCLUSIONS
High-purity oxygen aeration in the activated sludge process was further
evaluated during a four month contract study at the Batavia, New York
Municipal Pollution Control Plant. During the study, two distinct opera-
ting conditions were maintained. During the first phase of operation,
all of the wastewater flow for the City was treated in one of four avail-
able aeration bays using oxygen as the aerating gas. In the second and
final phase of operation, one-half of the wastewater flow was treated in
one bay using oxygen aeration while the other half was treated in two bays
with air aeration.
A main objective of this second contract was to evaluate in depth the
solids production and solids profile through the oxygenation system.
Solids and mass balances were performed on the system for each week of
the operation. The results of the mass balances indicate that a valid
accounting of the solids was obtained during the study. The solids
balances generally confirm the sludge production figures originally deter-
mined for the oxygenation system during the initial contract work at
Batavia. Therefore, it is concluded that the operating conditions main-
tained in the Batavia oxygenation system as compared with the operating
conditions maintained in Batavia's conventional plug flow diffused air
aeration system will yield 40-45% less excess biological sludge production.
A sludge production correlation was obtained which utilized data from
this contract as well as the data published from the initial contract
work. This correlation was used to prepare revised economic comparisons
of oxygenation and conventional diffused air aeration wastewater treat-
ment systems. These economic studies indicate that total treatment costs
in cents per thousand gallons treated with oxygenation are comparable
to costs with conventional diffused air aeration for a plant size of
1 MGD. As plant size increases, costs with oxygenation becomes less
than costs with conventional diffused air aeration with savings of
about 20% projected at this time for a 100 MGD design.
A second major objective of this contract was to investigate the filter-
ing characteristics of waste activated sludge from the Batavia oxygenation
system. To accomplish this, waste activated sludge from the clarifier
underflow was filtered on a pilot scale (3 ft diameter x 1 ft length)
vacuum filter. Vacuum filtration tests were made both with and without
prior thickening. From these tests, it is concluded that waste activated
sludge from an oxygenation system can be vacuum filtered directly without
thickening. The sludge cake yield obtained on unthickened sludge varied
from 1.0 Ib/hr/ft^ to 4.5 Ib/hr/ft2} depending on cycle time and solids
concentration in the feed sludge. Waste activated sludge from the oxygen-
ation system was also thickened up to 4.5% by simple gravity settling
prior to filtering. With thickening, the maximum sludge cake yield
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increased to over 5.0 lb/hr/ft2. Under the conditions tested, ferric
chloride was the most effective chemical conditioner for both thickened
and unthickened waste activated sludge. The optimum dosage of ferric
chloride was 200 pounds (FeCl-i) per ton of dry solids. At this dosage,
the pH of the conditioned sludge decreased to around 3.0. This pH value
could possibly be used to control the addition rate of ferric chloride
in a plant size operation. Comparison of available data on the vacuum
filtration of thickened waste activated sludge from one air aeration
system with data obtained in this contract indicates that oxygenation
system sludge will filter better and achieve a higher cake yield under
similar operating conditions.
A third major objective of this contract was to evaluate aerobic
digestion of oxygenation system waste activated sludge with high-purity
oxygen aeration. The aerobic digestion studies were performed entirely
on a batch basis. A non-objectionable, stabilized sludge was obtained
in 7-9 days with a 20-30% reduction in volatile suspended solids. The
stabilized sludge, under the conditions studied in this contract, was
more difficult to filter than the oxygenated waste activated sludge.
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SECTION II
RECOMMENDATIONS
This contract has demonstrated the ability of conventional sludge
handling processes to successfully handle oxygenated waste activated
sludge and in the case of vacuum filtration quite possibly to increase
the yield of the vacuum filter over that experienced with vacuum
filtration of conventional air aerated waste activated sludge. The
higher concentration of waste solids from an oxygenation system results
in a proportional decrease in the volume of sludge to be handled.
Further reductions in waste sludge quantity are also indicated because
of the anticipated decrease in pounds of dry solids produced by an
oxygenation system. All of these factors suggest that significant
potential savings may be achievable in sludge handling and disposal
operations when using oxygen in the activated sludge process. It is
recommended that additional sludge handling studies be undertaken on
a larger scale to verify the vacuum filtration data generated in this
contract and to investigate the performance of other sludge handling
processes on oxygenated sludges.
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SECTION III
INTRODUCTION
REVIEW OF 1969 OXYGENATION CONTRACT RESULTS
In 1969, Union Carbide Corporation undertook a full-scale evaluation of
high-purity oxygen aeration in the activated sludge treatment of munici-
pal wastewater. This work was conducted under FWPCA Contract Number
14-12-465 and has been reported in Water Pollution Research Series
Report Number 17050DNW05/70.l
During this initial contract work, a practical and economically attractive
means for using high-purity oxygen as the aerating gas in the activated
sludge process was evaluated over an eight-month period of continuous
operation. A temporary oxygenation system was installed in one half of
the existing aeration tanks at the Batavia, New York Municipal Pollution
Control Plant. The aeration tanks, which are of conventional diffused
air design, were covered and equipped with gas-liquid staging baffles to
provide a multistage system. Sparger impeller gas-liquid contacting units
were employed for oxygen transfer in each stage. The oxygenation system
was operated with co-current gas and liquid flow in the aeration (oxygena-
tion) tank.
The initial contract test program was specifically planned to evaluate the
following:
1. The feasibility of using high-purity oxygen for aeration in the
activated sludge waste treatment process.
2. The comparison of oxygenation and conventional diffused air aeration
in terms of overall treatment performance.
3. The consistent operation of the oxygenation system at high mixed-
liquor suspended solids (MLSS) levels (6000-8000 mg/1), high mixed-
liquor dissolved oxygen (D.O.) concentrations (8-10 mg/1), and high
overall utilization of feed oxygen gas (>907o) •
4. The operation and treatment performance of an oxygenation system at
low detention times (approximately 1.5 hours based on raw wastewator
flow) and high organic loading conditions considered economically
desirable but impractical with conventional diffused air aeration.
5. The relative economics of oxygenation comparer! with conventional
diffused air aeration.
The work was conducted in three phases. In two phases, the performance
of the oxygenation system was directly compared with that of the parallel
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conventional air aeration system, with each system treating half of the
plant's wastewater flow. In a third phase, the oxygenation system was
used to treat the entire plant wastewater flow in one-fourth of the plant
aeration tankage. A summary of results from each of the three phases of
operation is shown in Table 1.
The oxygenation system consistently exhibited superior treatment perfor-
mance in direct comparison to the parallel conventional air aeration
system, even though high aeration rates were used in the air aeration
system. The electrical power required for oxygen dissolution in the oxy-
genation system ranged from 0.08-0.14 drawn HP/1000 gallons of mixed-
liquor under aeration.
The biomass generated by the oxygenation system was highly flocculant and
rapidly settleable, yielding weekly average sludge volume indices (SVl's)
between 35 and 87 and settled sludge concentrations in the secondary
clarifier underflow of 2-3% (dry solids basis) with a wastewater tem-
perature that varied from a low of 53°F to a high of 73°F. The high
suspended solids concentration of waste activated sludge from the
oxygenation system can possibly obviate the requirement in many plants
for separate thickening of this waste stream prior to further sludge
processing, resulting in cost savings in the overall sludge handling and
disposal operation.
An extensive evaluation of oxygenation system total treatment costs in
comparison to conventional diffused air aeration total treatment costs
indicated significant savings could be achieved with the use of oxygen
aeration. These savings are due to several characteristics of the oxy-
genation system, but the major factors are reduced aeration tank volume
requirements for equivalent treatment, reduced power requirements, and a
reduced production and disposal cost of waste activated sludge. The
results of the economic evaluation indicate that the savings become
increasingly significant as the total quantity of wastewater treated is
increased. At a plant size of 1 MGD, air and oxygen aeration total
treatment costs were shown to be comparable. At 6 MGD, the use of
oxygen aeration afforded a 20 percent savings in total treatment costs.
At the larger plant sizes of 30 and 100 MGD considered in the study,
uotal treatment cost savings of about 25 percent were projected with
oxygenation. Total treatment costs referred to here include investment
and operating costs associated with preliminary treatment, primary treat-
ment, secondary treatment by the activated sludge process, and the dis-
posal of primary and waste activated sludges, all reduced to cents per
thousand gallons treated.
DISCUSSION OF SIGNIFICANT TECHNICAL RESULTS
FROM 1969 OXYCENAT10N CONTRACT
The prime objective of the initial contract was to demonstrate that the
uxygenation system could achieve increased treatment efficiency at high
solids levels and low detention times, while laaintaining high oxygen
utilization and low total power consumption. While demonstrating the
u
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TABLE 1
COMPARATIVE SUMMARY OF AIR AERATION AND OXYGENATION SYSTEM PERFORMANCE
Raw Wastewater Feed Rate
Sludge Recycle R
% Sludge Recycle
Raw Wastewater T
Mixed-Liquor Ave
Clarifier Efflue
*Aeration Detenti
**Nominal Aeration
MLSS Concentrati
MLVSS Concentrat
Recycle Sludge T
Recycle Sludge V
Recycle VSS/TSS Ratio
Mixed-Liquor S'
Mixed-Liquor Ii
(ft/hr)
***Final Clarifier Surface Overflow
Rate (gpd/ft )
1969 CONTRACT
Phase I
Operation
Air Aeration Oxygenation
System System
d Rate (MGD)
e {MGD)
iperature (°F)
ge D.O. Cone . (mg/1)
D.O. Cone. (mg/1)
i Time (hr)
•etention Time (hr)
i (mg/1)
m (mg/1)
(mg/1)
1 (mg/1)
tio
;e Volume Index
.al Settling Velocity,
1.97
0.25
13.0
59
1.5
0.4
3.5
4.0
2440
1740
15,000
9,700
0.65
76
7.7
1.91
0.44
24.0
59
8.7
4.0
3.4
4.1
3060
2210
18,600
11,800
0.63
64
7.2
Phase II Operation
Oxygenation
System
2.53
0.83
34.0
66
9.0
0.7
1.2
1.5
6980
4450
29,600
18,200
0.62
36
6.5
Phase III
Air Aeration
System
1.29
0.25
21.0
70
O.S
0
2.6
3.0
3640
2580
16,600
11,800
0.71
63
7.7
Operation
Oxygenation
System
1.44
0.60
45.0
70
8.0
4.3
2.0
2.8
6190
4310
18,800
12,700
0.69
49
5.1
1570
1520
1010
1030
1140
* Raw flow plus recycle sludge flow.
** Raw wastewater flow only.
*** Based on total clarifier surface
area of 1260 ft2.
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TABLE 1 CONTD.
COMPARATIVE SUMMARY OF AIR AERATION OXYGENATION SYSTEM PERFORMANCE
1969 CONTRACT
Raw wastewater BOD (mg/1)
Clarifier Effluent BOD (mg/1)
7, BOD Removal
Raw Wastewater COD (mg/1)
Clarifier Effluent COD (mg/1)
% COD Removal
Raw Wastewater TSS (mg/1)
Clarifier Effluent TSS (mg/1)
% TSS Remova1
Raw Wastewater VSS (mg/1)
Effluent Turbidity (J.T.U.)
Food/Biomass Loading
(Ib BOD applied/day/lb MLVSS)
Volumetric Organic Loading
(Ib BOD applied/day/1000 ft mixed-liquor)
Lb Dry Solids Wasted/Day
Lb VSS Was ted/Lb BOD Removed
Phase I
Operation
Air Aeration Oxygenation
System System
159
16
90
352
84
76
221
16
93
152
5.7
0.57
159
11
92
352
73
80
221
9
96
152
5.6
0.42
Phase II Operation
Oxygena tion
System
220
23
90
325
97
71
174
19
89
123
9.1
0.79
Phase III
Air Aeration
System
262
30
88
578
116
79
430
23
94
332
5.7
0.84
Operation
Oxygenation
System
262
14
94
578
89
84
430
12
97
332
4.4
0.55
60.0
3150
0.87
57.9
1800
0.48
212.5
2450
0.41
128.9
3500
0.99
144.8
590
0.13
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TABLE 1 CONTD.
COMPARATIVE SUMMARY OF AIR AERATION AND OXYGENATION SYSTEM PERFORMANCE
1969 CONTRACT
Phase I Operation Phase II Operation
Raw Wastewater Total Phosphorus (mg/1)
Clarifier Effluent Total Phosphorus (mg/1)
7» Total Phosphorus Removed
Raw Wastewater N^-N as N (mg/1)
Clarifier Effluent NH3-N as N (mg/1)
Raw Wastewater TKN (mg/1)
Clarifier Effluent TKN (mg/1)
Raw Wastewater N02~N as N (mg/1)
Clarifier Effluent N02-N as N (mg/1)
Raw Wastewater N03-N as N (mg/1)
Clarifier Effluent N03~N as N (mg/1)
Raw Wastewater Total Nitrogen (mg/1)
Clarifier Effluent Total Nitrogen (mg/1)
% Total Nitrogen Removed
N03-Xitrogen/Total Nitrogen in Clarifier
Effluent
Lb Total Nitrogen Removed per
Lb Dry Solids Wasted
Air Aeration
System
11
6
44
20
17
34
22
0.2
0.4
0.3
0.6
35
23
35
0.03
0.066
Oxyge nation
System
11
7
39
20
15
34
19
0.2
0.9
0.3
1.6
35
21
37
0.08
0.129
Oxygenation
Sys tern
12
10
12
19
18
33
22
0
0.2
0.1
1.7
34
26
23
0.07
0.066
Phase III Operation
Air Aeration
System
19
13
32
30
24
53
31
0
0.2
0.1
2.2
53
33
36
0.07
0.060
Oxygenation
System
19
14
27
30
24
53
30
0
0.7
0.1
2.5
53
33
36
0.07
0.453
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TABLE 1 CONTD.
COMPARATIVE SUMMARY OF AIR AERATION OXYGENATION SYSTEM PERFORMANCE
Ft Air Utilized/
Feed Oxygen Gas F
Gas Exhaust Rate
Final Stage (cfm.NTP)
Oxygen System Exhaust Gas
Composition (% Oxygen)
Mixing (Shaft HP)
Overall Power Utilized
1000 Gallons mixed-liquor)
Total Air Blower Power Utilized
Ion Wastewater Treated
• Rate (cfm.NTP)
>m Oxygen System
NTP)
t Gas
v pen)
J b '
ygen Utilized
Lb BOD Consumed
for Liquid
for Gas
70% Eff .) (HP)
ed (HP/
1969 CONTRACT
Phase I Operation
Air Aeration Oxygenation
System System
2.89
18.3
1.74
46
95.5
0.94
26.0
2.6
0.088
Phase II Operation
Oxygenation
System
_
31.0
4.04
55
92.7
0.96
11.1
9.2
0.125
Phase III Operation
Air Aeration
System
4.32
Oxygenation
System
19.0
3.22
51
91.4
0.72
11.7
10.6
0.137
(HP)
141
138.5
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above, it was also confirmed that the oxygenation system produced a
smaller mass of waste activated sludge than a conventional air aeration
system, and that this sludge concentrated to 2-3% in the clarifier under-
flow. These factors suggested that further economics are possible in
wastewater treatment with oxygenation, not only because fewer pounds of
excess secondary sludge are produced for subsequent processing, but also
in the manner in which the secondary sludge is handled. For these
reasons, a second contract was undertaken to further evaluate oxygena-
tion system kinetics and sludge production, and to investigate on a
pilot scale the aerobic digestion and vacuum filtration characteristics
of the system's waste sludge.
The following conclusions were made from the initial contract work con-
cerning sludge settling, sludge production, and sludge handling:
Sludge Settling
Oxygenated biomass settles more rapidly than typical air aerated acti-
vated sludge biomass. Although initial settling rate and compactibility
advantages (at the same MLSS concentration) exist over the common range
of mixed-liquor suspended solids operating levels used in typical air
aeration systems (1500-3500 mg/1), they are more apparent when the oxy-
genation system is operated at high solids levels as demonstrated during
Phase II of the initial contract (6000-8000 mg/1). It is significant
that oxygenated biomass at a mixed-liquor concentration in excess of
6000 mg/1 achieves an initial settling rate comparable to a typical air
system biomass operating at a mixed-liquor concentration of 2000-
3000 mg/1. If this were not the case, a system designed for these high
solids levels would be unable to achieve effective clarification and
thickening without significant additional expenditure in clarifier
capacity.
The excellent dewatering and compactibility characteristics of oxy-
genated biomass also are important in clarifier design. At the high
mixed-liquor solids concentrations typically carried in an oxygenation
system, the mass loadings imposed on the clarifier are significantly
higher than for typical air aeration system operation. Whereas an
air aeration system clarifier may be operated at a mass loading of 15
to 25 pounds TSS/day/ft2, oxygenation system clarifiers generally
operate at mass loadings of 45-65 pounds TSS/day/ft2- The excellent
compactibility of oxygenated sludge is the factor which permits con-
ventionally sized clarifier designs to handle these magnitudes of mass
loading effectively and without loss of clarification or thickening
capability.
In Phase II of the initial contract, high mixed-liquor solids were
maintained and these solids settled and compacted well. While an air
aeration system could probably be designed to maintain a high D.O. at
an MLSS in excess of 6500 mg/1, the energy input would be uneconomical
and would contribute to very high floe shear and consequent destruction
11
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of the sludge flocculating properties, so that settling and solid-liquid
separation efficiency would be significantly impaired.
Thus, the use of oxygen aeration has the dual effect of both allowing high
solids levels to be carried practically and economically, and at the same
time achieving high mixed-liquor D.O. without excessively shearing the
mixed-liquor biomass.
Sludge Production
At comparable biomass loadings, a significantly lower net sludge produc-
tion is achieved with an oxygenation system in comparison with a conven-
tional air aeration system. The fundamental reasons which have been
proposed for the decreased sludge production, and thus longer solids
retention times, in an oxygenation system include staged operation at
higher MLSS and D.O. levels with consequently higher auto-oxidation
continually taking place and alteration of metabolic pathways due to
such effects as increased D.O., lower pH, and operation at high oxygen
partial pressures.
Recent additional confirmation of lower sludge production with an oxygena-
tion system at comparable biomass loadings can be found in the literature.
Poon^ made the following observation about the solids production as a
function of substrate removal. "It is recognized that solids production
as percentage of substrate removed increases with process loading. The
conventional activated sludge system at a process loading of approximately
0.5 would produce an average sludge synthesis of 45 percent of the weight
of COD removed. When process loading is increased to 3.0, sludge produc-
tion would increase to 55-60 percent. This is particularly true for high
solids systems where insufficient dissolved oxygen may slow down the
endogenous respiration resulting in a buildup of non-viable cells.
Rickard et al reported that the amount of carbohydrate production of
cells in aeration basins decreased as the D.O. concentration was increased,
The variation in carbohydrate production could be a direct effect of
oxygen tension in the metabolic fate of the substrate. The fact that the
protein and RNA content of cells did not change in Rickard's study^^
indicates the possibility of higher endogenous respiration rates in the
high oxygen tension system which oxidized more carbohydrate material."
The rate of solids formation calculated from some of Poon's data is
shown plotted against the rate of COD removal in Figure 1. The figure
shows that high dissolved oxygen concentrations maintained by oxygen
aeration resulted in lower solids production than did the lower dis-
solved oxygen concentrations that air aeration was able to maintain.
Poon3 concludes "High oxygen tension maintained by pure oxygen aeration
can provide a high efficiency of COD removal in a system of high solids
^nd for a process loading range between 0.7 to 4.0 gm COD per gm SS per
Jay ...Smaller amounts of sludge produced per unit weight of COD removed
adds to the benefit of such a system."
Union Carbide has basic as well as applied development programs in
12
-------
FIGURE I
SOLIDS PRODUCTION
CALCULATED FROM ROOM'S DATA3
SOLIDS •< 4,800 TO 10,900 MG/L
F/M RATIO-0.7 TO 4.0 LB COD/DAY/INITiAL LB SOLIDS
2.4r
CO
9
3 2.0
03
_l
_l
<
t 1.6
x^
O
Nw
FORMED,
K>
9
i °-8
CD
0.4
r»
0
AIR AERATION SYSTEM
& DO §6 MG/L
•
o
0
0 °
0
0
OXYGEN AERATION SYSTEM
O ° o DO^ 15 MG/L
O
i i i i i
012345
LB COD REMOVED /(DAY/ INITIAL LB SOLIDS
13
-------
progress, designed to further quantify the underlying basic factors
governing the sludge production and sludge settling characteristics of
oxygenation systems. Extensive pilot planting has been and still is
being conducted in Union Carbide owned pilot plants in many different
locations in the United States. These pilot plants handle a variety of
feed streams encompassing both unsettled raw wastewater and primary
effluent from both municipal and industrial sources. The attached
Table 2 includes a summary of the results obtained with some pilot
plants .
Sludge Handling
The settling and compacting characteristics of oxygenated sludge result in
a consistently high clarifier underflow solids concentration, permitting
relatively low sludge recycle rates with high aeration tank MLSS levels
and lower volume transfer in sludge wasting operation. Further advantages
will accrue in subsequent sludge disposal processes.
One major process in use in wastewater treatment plants is vacuum filtra-
tion. This operation is commonly conducted on raw primary sludge and
anaerobically digested sludge. The direct filtration of waste activated
sludge is not widely practiced because of the general inability of vacuum
filters to produce a cake in yields high enough to be economic. Economics
favor either recycling to primaries, anaerobic digestion, or other means
of conditioning to make the waste sludge filter more easily. In a few
cases, waste activated sludge is thickened separately and then vacuum
filtered.
The increased solids concentration obtained in the initial contract work
and the observed excellent flocculating characteristics of the oxygenated
biomass would indicate that direct vacuum filtration of waste activated
sludge might be practical and economical. Vacuum filtration of waste
activated sludge represents one of the major objectives of this contract.
A second solids disposal process receiving increasing attention in waste
treatment technology is aerobic digestion. This operation achieves a
decrease in total solids and produces a stable product in a shorter time
than is required by anaerobic digestion. More importantly, aerobic diges-
tion does not have the following disadvantages inherent with anaerobic
digestion.
1. High supernatant BOD and suspended solids.
2. High initial cost.
3. Need for higher feed solids concentrations.
4. Many operational problems.
Aerobic digestion of oxygenated sludge will result in digester volume
-------
TABLE 2
STOIARY OF DATA FROM TO I OX CARBIDE PILOT PLANT PROGRAMSv
Reactor Detention line (hr)—
MLSS (mg/1)
MLVSS (mg/1)
Flow Rate (gpm)
Food/Biomass Loading (lb BOD Applied/day/
lb MLVSS)
Volumetric Organic Loading (lb BOD Applied/
day/1000 ft3)
Average Wastewater Temperature (°F)
Cl=rifier Surface Overflow Rate (gpd/ft2)
Clarifier Mass Loading (lb TSS/day/ft2)
Recycle Sludge (•". dry solids)
Wastewater Type
Wastewater Pre-Treatoent
Potential Full Scale Plant Flow (mgd)
Wastewater Feed to Oxygenacion System COD
(mg/1)
Removal Across Oxyse-acion System (7.) COD
Lb VSS Forred lr JOD Re^.uved
Lb 02 L'tili:. = c '.? 300 Rer.cved
I
1.3
5950
4760
20.5
id/day/ 0.64
Applied/ 196
86
•d/ft*) 890
'ft2) 52
2.3
Municipal
Degritted
;d) 80
:em COD 398
BOD 170
SS 172
%) COD 84
BOD 97
SS 92
0.36
1.07
II
1.8
4700
3950
1.9
1.4
350
65
520
31
1.5
Municipal
Grain Process
Meat Packing
Primary
30
826
415
180
88
97
90
0.63
0.9
III
2.0
4750
3850
13.5
0.81
191
68
585
32
2.3
Municipal (407.)
Industrial(60%)
Primary
150
652
253
171
66
92
80
0.60
1.09
IV
1.8
6100
3800
15.0
0.66
165
57
660
45
2.6
V
3.5
5850
4360
7.6
0.54
147
83
330
20
3.2
VI
1.8
7350
5270
15.0
0.6
197
78
650
50
3.2
Municipal (507,) Municipal Municipal
Industrial(507,) Textile(607.)
Primary
105
375
185
180
77
90
89
0.71
0.89
Degritted
8
690
343
317
82
91
93
0.30
1.16
Degritted
110
433
237
282
85
95
90
0.36
1.0
VII
2.6
5100
4550
10.3
0.6
170
65
900
37
1.6
Municipal (50%)
Industrial (507=,)
Primary
UO
930
295
110
70
91
63
0.69
1.66
More detailed information on these pilot plant studies is available by writing to
Vr.ion Carbide Corporation, Linde Division, P. 0. Box 44, Tonawanda, N. Y. 14150,
C 0 Marketing Department
Based on raw Wastewater flow only.
-------
savings due to the higher initial solids concentration of the waste sludge
feed to the digester. The use of oxygen aeration in an aerobic digester
may provide further economic and process advantages. The study of aerobic
digestion of oxygenated sludge (using oxygen aeration) was another major
objective of this contract.
16
-------
SECTION IV
CONTRACT PROGRAM PLAN
The operation of the oxygenation system at Batavia was continued in
1970 to further delineate system kinetics and to investigate the
dewatering and stabilization characteristics of oxygenated waste activa-
ted sludge. The specific objectives of the 1970 contract were to
evaluate:
1. The biokinetic characteristics of the system as revealed by
intensive analysis of the system with heavy emphasis on the
solids profile and sludge production.
2. The sludge dewatering aspects of waste activated sludge from the
full scale oxygenation system by direct vacuum filtration using
a pilot scale filter (3' diameter x 1' length).
3. The aerobic stabilization (digestion) of oxygenation system waste
activated sludge using a pilot scale aerobic digester (oxygen
aerated) followed by vacuum filtering of the digested sludge using
the pilot scale filter.
The full scale oxygenation system was operated for a period of over four
months during the summer and fall of 1970. Very few equipment changes
were required as the system had been operated extensively during the
original 1969 contract work. The oxygenation system was re-started
during late July: checked out and refined during August, and was ready
to begin complete data acquisition beginning in September.
The 1970 contract was performed in two phases. In the first phase of
operation (designated as Phase I) during September and October, all of
the Batavia wastewater flow was diverted through one bay of the Batavia
facility in which a three-stage oxygenation system was employed for
aeration. In Phase II during November, the wastewater flow was split
equally between the oxygenation system and the air aeration system.
Data acquisition was limited to the oxygenation system during this
period.
During both phases of operation, waste sludge obtained directly from the
clarifier underflow of the oxygenation system was subjected to aerobic
digestion and vacuum filtration studies. The vacuum filtration studies
were performed using a pilot-scale Eimcobelt vacuum filter. Anaerobically
digested sludge and aerobically stabilized sludge were also filtered
using this apparatus.
Aerobic stabilization studies using waste activated sludge from the
oxygenation system and oxygen as the aerating gas were carried out in
two pilot-scale aerobic digesters. These studies were conducted as
batch experiments in which the sludge was charged to the digester and
and allowed to digest in an oxygen atmosphere for varying periods of
time. The digested sludge was then filtered using the pilot-scale
vacuum filter.
17
-------
SECTION V
EQUIPMENT DESCRIPTION AND EVALUATION METHODS
MODIFICATIONS TO OXYGENATION SYSTEM
At the completion of the initial FWPCA contract work in 1969, the
oxygenation system was mothballed and the entire operation resumed using
air aeration. Very few changes were necessary, therefore, to adapt the
oxygenation equipment for operation to meet the present contract objec-
tives. Appendix A of this report contains considerable detail pertinent
to the modifications made to the Batavia plant for the initial contract
work. The discussion below is only pertinent to additional modifications
made for the 1970 contract.
The most significant change made at the site was the installation of two
variable speed centrifugal sludge return pumps. These 1 MGD capacity
pumps with General Electric Model VT700 SCR controls were installed so
that the return sludge rate could be more accurately controlled. Each
return pump was equipped with a remote control station so that operation
could be continuously monitored and easily adjusted. The installation
of these pumps alleviated the control problem which had occurred during
the initial contract, particularly during Phase III.
The compressors that were used in the initial contract performed very
satisfactorily during the entire period of that contract and would have
been satisfactory for use in the present contract. These compressors,
however, were selected to serve only a one to two year period of service
consistent with the initial contract objectives. Since Union Carbide
was interested in evaluating compressors designed for many years of
operation, it was decided to purchase and install new compressors at
Union Carbide expense. The gas blowers used for oxygen gas recirculation
in the three stages were low speed rotary lobe type. The Union Carbide
sponsored evaluation was undertaken to determine suitability of these
compressors on the basis of design, materials, maintenance, reliability,
oxygen safety, and overall ability to function in this type of service.
Approximately one month prior to the planned start-up date, work was
begun to completely check-out the oxygenation equipment and prepare it
for return to service. All of the mixers were checked, oiled, and run
for short periods. All instrumentation at the aeration tanks and in the
control room was checked and calibrated.
A pilot-scale 3 ft diameter Eimco rotary belt vacuum filter was rented
for the duration of the contract. This unit was installed in a building
located adjacent to the aeration tanks and all necessary piping connections
and auxiliary services tied in so that the vacuum filtration studies
with oxygenated waste activated sludge could be conducted. Additional
piping was installed so that sludge could be obtained from the plant
anaerobic digesters and the pilot-scale aerobic digesters.
19
-------
The oxygen aeration pilot plant used in the initial contract studies was
modified for conducting aerobic digestion studies using oxygen. Two 800-
gallon batch aerobic digesters were made by partitioning the 1600-gallon
pilot plant into two separate tanks. Each digester was equipped with
oxygenation and mixing devices. All instrumentation necessary to monitor
the operation of the digestion cycles was installed. This equipment was
set up to be operated on a fill and draw basis in which varying aerobic
digestion periods could be established and monitored. The digested sludge
was then to be pumped to the Eimcobelt filter for determination of the
filtering characteristics of aerobically digested sludge.
DATA ACQUISITION AND ANALYSIS
As was done in the initial contract at Batavia, extensive sampling and
analytical procedures were established to obtain performance information
on the oxygenation system. All of the basic information obtained in the
initial contract, as well as some additional data, were taken. A summary1
of the data acquisition with specific reference to the additional data
collected is as follows:
1. Raw wastewater, sludge recycle, and waste activated sludge flows for
the oxygenation system including streams recycled to the oxygenation
system.
2. Total and volatile suspended and total and volatile dissolved solids
concentrations in the raw wastewater, recycle sludge, mixed-liquor,
and clarifier effluents for the oxygenation system. (Dissolved
solids analyses were not performed on the initial contract.) Solids
analyses were also performed on all of the recycle streams to the
oxygenation system.
3. Physical characteristics of the mixed-liquor as determined by sludge
volume index (SVI) and settling velocity measurements in the
oxygenation system.
4. Chemical and biochemical characteristics of the raw wastewater,
mixed-liquor, and final settler effluent as revealed by BOD, COD,
ammonia nitrogen (NH3-N), nitrite nitrogen (N02-N), nitrate nitrogen
(N03-N), total Kjeldahl nitrogen (TKN), and total (TP) and ortho-(OP)
phosphorous determinations in each case. Both total and soluble
samples were analyzed on this contract.
5. Dissolved oxygen concentrations at several points in the oxygenation
system were continuously monitored and recorded.
6. Gas-phase oxygen composition in the gas space of each of the stages
in the oxygenation system and in the vent gas was continuously mon-
itored and recorded.
7. Overall oxygen utilization in the oxygenation system by measurement
20
-------
of feed and exhaust gas volumes and exhaust gas oxygen composition.
8. Measurement of actual power consumed for gas recirculation and
liquid mixing in the oxygenation system.
9. Measurement of all temperatures and pressures necessary in various
parametric calculations.
10. Measurement of gas recirculation rates within stages of the
oxygenation system.
Continuous flow-weighted composited sampling of influent raw wastewater
and final settler effluents was carried out throughout the course of the
work. These 24-hour composite samples were analyzed daily for BOD, COD,
solids, phosphorous, and the nitrogen series. Both total and dissolved
analyses were performed.
Grab samples of mixed-liquor and recycle sludge were taken twice daily
and separately composited for the Total Suspended Solids (TSS), Vola-
tile Suspended Solids (VSS), Total Dissolved Solids (TDS), and Volatile
Dissolved Solids (VDS) analyses.
Where SVI or settling velocity of the mixed-liquor activated sludge was
to be measured, separate grab samples were taken from each aeration stage
and evaluated immediately. Supernatant fractions of these samples
(unfiltered and filtered) were analyzed for BOD, COD, NI^-N, N02~N,
N03-N, TKN, and TP content periodically.
When sludge was wasted from the oxygenation system (usually a 4-5 hour
period/day), separate grab samples of recycle sludge were taken during
this period for VSS and TSS determinations. This permitted a more
accurate estimation of the quantity of dry solids wasted from the system.
Union Carbide technicians were on-site only 8 hours per day for five
days each week to conduct analyses and monitor the systems. During
weekends and evenings, off-duty plant operators and temporary help were
retained to collect samples, conduct certain analyses, and monitor the
operation. No regular plant operating personnel or Union Carbide
technicians were on duty from 11:00 PM each night to 7:00 AM each
morning. This was a period of unattended operation. Late night samples
were taken by off-duty operators or part-time temporary help just before
leaving the plant at 11:00 PM.
The methods used to make individual measurements and laboratory deter-
minations were the same as used in the initial contract. A summary of
these methods is provided in Appendix B of this report.
DATA REDUCTION
In day-to-day operation throughout the course of the work, only raw data
21
-------
from analysis and monitoring were collected at the work site. In weekly
batches, these data were computationally reduced at Union Carbide's
Tonawanda facility using an IBM 360/50 computer and a program specifically
written to analyze and reduce the raw data to meaningful form. The product
of this approach to data reduction was a listing of all daily and weekly
system operating conditions and performance results. This means of data
reduction was particularly advantageous in that it permitted significantly
more observations and analyses to be made since technicians were relieved
of devoting time to calculating and cataloging results. The quantity of
data obtained for each day of operation are so extensive that they are
not listed on a daily basis in this report. Weekly averages only are
listed in Appendices C through E.
SOLIDS AND MASS BALANCE METHODS
Solids and mass balances were performed on the oxygenation system for
each week of operation during the contract. The results of these balances
are discussed in detail later in this report. To perform these balances,
it was necessary to analyze all recycle stream to the oxygenation system
and to incorporate the results of these analyses into a consistent solids
and mass balance calculation procedure. The calculation procedure used
is outlined below.
1. Calculate the average metered flow of wastewater to the oxygenation
tanks (MGD) from daily integrator readings over a weekly period.
2. Calculate the average waste sludge flow (MGD) from the daily sludge
wasting log.
3. The TSS concentration of thickened waste sludge was found to average
about 39,000 mg/1 for a number of samples analyzed during the
contract period.
4. Calculate the average thickened waste sludge flow (MGD) = average
waste sludge flow rate x mg/1 TSS in waste activated sludge.
39,000
5. The average thickener supernatant flow (MGD) was equal to 0.15 x
average thickened waste sludge flow.
6. Calculate the average digester supernatant flow (MGD) from integrator
readings at the beginning and at the end of the week.
7. Calculate the average vacuum filtrate flow (MGD) = average thickened
waste sludge flow - average digester supernatant flow.
8. The average chlorine contact tank transfer was 0.00714 MGD (50,000
galIons/week) .
9. Calculate the average corrected raw wastewater flow rate (MGD) =
22
-------
average metered flow to the oxygenation tanks - average vacuum
filtrate flow - average chlorine contact tank transfer - average
thickener supernatant flow.
10. Calculate the average corrected clarifier effluent flow (MGD) =
average metered flow to the oxygenation tanks - average waste
sludge flow + average digester supernatant flow.
11. Calculate the average Ib/day of TSS, VSS, IDS, and VDS contained
in the raw wastewater, vacuum filtrate, digester supernatant,
thickened waste sludge, and clarifier effluent streams from the
flows and appropriate solids concentrations.
12. It was estimated that a layer approximately 6" deep collects on
the bottom of the chlorine contact tanks each week which has a TSS
concentration of 30,000 mg/1 and a VSS concentration of 5,000 mg/1.
13. Calculate the average Ib/day of TSS, VSS, TDS, and VDS returned to
the system by transfer from the chlorine contact tanks from (8) and
(12).
14. Calculate the average Ib/day of TS, VS, and IS in all streams.
TS = TSS + TDS
VS = VSS + VDS
IS = TS - VS
The weekly solids balances prepared for both Phase I and II operation
of the oxygenation system are presented later in this report.
The mass balances were made over the portion of the system enclosed by
the broken line in Figure 2. A mass balance is made by correcting the
total solids balance to account for oxygen, carbon, and hydrogen which
enters or leaves the system in gaseous or liquid form and which is not
measured in the determination of total solids.
By definition, let:
Mass In = Total Solids In + delta 02 in feed 02
Mass Out = Total Solids Out + delta C in waste gas
+ delta 02 in waste gas
+ delta C in clarifier effluent
+ delta 02 in clarifier effluent
+ delta H2 in clarifier effluent
23
-------
The additional calculations used to estimate the various delta values
and to convert a total solids balance to a mass balance follow:
15. The weekly average value of delta 62 in feed C^ (lb 02/day is the
same as the weekly average Ib/day of feed oxygen.
16. Calculate the average mol fraction of carbon dioxide in the waste
gas (YC02):
YC02
3
where T = alkalinity, equivalents/liter
K^ = first ionization constant for carbonic acid
K2 = second ionization constant for carbonic acid
|H+j., = average hydrogen ion concentration of mixed
liquor in last stage, gm/liter
H_o = Henry's law constant for carbon dioxide,
2 gm mols/liter x atm.
17. Calculate the weekly average Ib/day of undissolved oxygen in the
waste gas stream from the measured flow and mol fraction (YC^)
values.
18. Calculate the weekly average value of delta C in waste gas
(lb C/day):
delta C in waste gas = lb/daV of Dissolved 0? x ^2. x u
32 Y02
19. Calculate the weekly average value of delta 02 in waste gas (lb
02/day):
weekly average lb/
delta 02 in waste gas = delta C in waste gas x — + day of molecular
C2 in waste gas
20. Calculate the average quantity of dissolved carbon dioxide leaving
the system dissolved in the clarifier effluent:
24
-------
RAW
WASTEWATER
r
SAMPLE
TAP
FIGURE 2
FLOW DIAGRAM FOR
BATAVIA OXYGENATION SYSTEM
MASS BALANCE ANALYSIS
CL2 CONTACT TANK TRANSFER
FEED 02 WASTE GAS
I
I
IFLOWMETER
\
OXYGENATION TANKS
CLARIFIER
K)
^RECYCLE SLUDGE
•FLOWMETER
'SAMPLE
-TAPS
WASTE
SLUDGE
THICKENER SUPERNATANT
DIGESTED
SLUDGE
VACUUM FILTER
VACUUM
FILTRATE
t
FILTER CAKE
THICKENER
O.2
CONTACT
TANKS
T
FINAL
EFFLUENT
CLEANING
OF TANKS
THICKENER
WASH WATER
,THICKENED WASTE SLUDGE
IDIGESTER NO. I
IDIGESTER NO. 2
DIGESTER SUPERNATANT
-------
Xyn moU CO? Ln GE = 43.6 x MGD of CE x T x
sec I Kj 2K]K2
+ K1K2
where [n Ip- = average hydrogen ion concentration of clarifier
*• •* effluent, gm/liter
21. The pH value of the clarifier effluent returns to the pH value of
the raw wastewater upon coming to equilibrium with air.
22. Calculate the average quantity of dissolved carbon dioxide
remaining in the clarifier effluent which has come to equilibrium
with air:
gn mols C02 at equilibrium = 43 8 x MGD of CE x T x/ rjpfl -
sec ] L J
\ *
\e
where H I = average hydrogen ion concentration of raw wastewater
L JR gm/liter
23. Calculate the weekly average quantity of excess carbon dioxide
leaving the system in the clarifier effluent:
lb mols excess CC^ _ 190.5 [gm mols CC-2 in CE _
day \ day
gm mols C02 at equilibrium\
day /
24. Calculate the weekly average values of delta C in clarifier effluent
(Ibs C/day):
delta C in clarifier effluent = 12 x lb m°ls excess C0?
day
25. Assume that cell mass can be represented by the formula
CH1.589°0.307N0.282S0.010'
26. Calculate the weekly average value of delta H2 in clarifier effluent
(lb H2/day) by estimating the quantity of hydrogen contained in
26
-------
the cell mass which is oxidized to water:
delta H2 in clarifier effluent = lj589 (delta C in waste gas + delta
12 C in clarifier effluent)
Note; Hydrogen may also be oxidized from the organic material in
the raw wastewater during the synthesis reactions. Allowance
for this oxidation could be made by changing the multiplier
1.589 in the above equation.
27- Estimate the net quantity of oxygen necessary to oxidize the weekly
average delta H2 in clarifier effluent (Ib combined 02/day):
2 _
H1.589°0.307 X 1.589 = H2.0°0.3864
Ib combined 0? = 16 (1-0.3864) x delta „ in clarifier effluent
day 2
28. Calculate the weekly average quantity of oxygen leaving the system
as dissolved oxygen in the clarifier effluent (Ib D.O./day):
V? D'°' = MGD of CE x 8.34 x mg/1 D.O. in CE
day
29. Calculate the weekly average value of delta 02 in clarifier effluent
(Ib 02/day):
delta 02 in _ Ib D.O. + Ib combined 0? +
clarifier effluent day day
/Ib mols excess CO? \
32 \ day /
30. Calculate mass in and mass out:
Mass In = Total Solids In + delta 02 in feed 02
Mass Out = Total Solids' Out + delta C in waste gas
+ delta 02 in waste gas
+ delta C in clarifier effluent
+ delta 02 in clarifier effluent
+ delta H2 in clarifier effluent
The weekly mass balances prepared for both Phase I and Phase II operation
of the oxygenation system are presented later in this report.
27
-------
SECTION VI
SOLIDS PROFILES AND EXCESS SLUDGE PRODUCTION
A prime objective of this contract was to investigate the non-volatile
solids profile as well as sludge production and wasting patterns of the
oxygenation system. To accomplish this objective, the oxygenation system
was operated in two distinct modes designated as Phase I and Phase II.
During each of these operating periods, extensive suspended and dissolved
solids analyses were conducted on all of the primary streams entering and
leaving the oxygenation system. Particular care was taken to insure that
flow weighted 24-hour composite samples were obtained. Additionally, it
was necessary to analyze the external recycle streams to the aeration
tanks from sources such as the vacuum filter, the anaerobic digester, and
the chlorine contact tanks, and to account for the gas streams and soluble
gases entering and leaving the aeration tanks.
This data acquisition program was established to insure thjt sufficient
flow weighted data was obtained from all entering and leaving streams to
permit an accurate mass balance to be made on the oxygenation system. If
j mass balance for the plant could be established, the excess sludge pro-
duction figures developed for the operation would be considered accurate
and representative. A mass balance has not commonly been used in analysis
of the activated sludge process. It differs from the solids balance
commonly employed in that a full accounting of oxygen, carbon, and
hydrogen entering and leaving the" system is made. It will be shown later
in this report that this accounting is valuable as a means of checking
and confirming the overall validity of analytical and sampling procedures
and as an aid to understanding of the process. A mass balance is not
generally required to establish solids production data except as a check
on procedures. In fact, it is quite difficult to make with an air acti
vated sludge process where all gas leaving the tanks is not contained and
vented at a single place.
OPERATING CONDITIONS - PHASE I
Data acquisition for Phase I operation was begun on August 30, 1970, and
was continued for eight full weeks through October 25, 1970. During Lhis
period, the entire raw wastewater flo^ to the plant was processed through
one-quarter of the available aeration tankage at the treatment plant.
Actual operation under this flow scheme was started prior to August 30;
however, data acquisition was not started until it was determined that
stabilized operating conditions had been established.
An overall diagram of the manner in which the Batavia plant was operated
during the Phase I period of September and October, 1970, is shown in
Figure 3. It can be seen that both aeration bays and one of the oxygen-
ation bays were shut down during this period. The total raw wastewater
flew was processed through bay 4 of the treatment plant. The mixed-
29
-------
OXYGEN SYSTEM
AIR SYSTEM
t
2.79 MGD
RAW
WASTEWATER
(AVERAGE FOR
PHASE I )
1 |
t
T
t
i
4
1
NOT
IN
SERVICE
BAY 3
UJ
0
o
in
z
-------
liquor from bay 4 was split and diverted to both clarifiers. Recycle
sludge was returned from both clarifiers to the front end of oxygenation
bay 4.
The operating conditions of Phase I were selected to provide sludge pro-
duction data at higher average food-to-biomass loadings than were exper-
ienced during the initial contract. The average food-to-biomass loading
during all of Phase I was about 1.0 pound of BOD applied/day/pound of
MLVSS. By comparison, the average loadings maintained in the initial
contract were:
Phase Lb BOD Applied/Day/Lb MLVSS
I 0.41
II 0.79
III 0.55
During the Phase I period of operation in 1970, no mechanical problems
occurred to affect the continuous operation of the plant. As was
repeatedly observed in the initial contract operation, the oxygenation
equipment continued to operate reliably with a minimum of operator
attention. All control systems and monitoring equipment functioned
normally during the operating period.
Table 3 on page 32 presents a summary of average operating conditions for
the Phase I operation. Appendix C of this report contains weekly summaries
of the operating conditions and results achieved during Phase I of the
contract.
The hydraulic data obtained indicates that the average raw wastewater
flow rate to the plant exceeded the plant design flow rate of 2.5 MGD
during all but one week of operation. Flow variations during this period
were typical of those experienced during the initial contract. While
daily flow would typically average 2.8 MGD, the hydraulic peak would
regularly reach a flow rate of about 3.5 to 4.0 MGD at approximately
10:00 AM and last for a duration of 5 to 8 hours. The minimum hydraulic
flow would occur during night-time hours at a flow rate of 0.8 to 1.0 MGD.
Flow weighted composite samples were taken throughout the contract work
to account for these flow variations. Recycle sludge flows were generally
maintained at 35 to 50 percent of the raw wastewater flow. Some trouble
was experienced in balancing the recycle flows from each clarifier.
However, careful attention to the sludge blanket level in each clarifier
helped to alleviate the problem. Average oxygenation times in the
aeration bay varied from 1.2 to 1.6 hours during this phase based on the
average raw wastewater flow rate.
The oxygenation equipment and dissolved, oxygen control equipment func-
tioned normally during the entire period maintained average dissolved
oxygen levels of 6-10 mg/1 (range 4.7 to 12.5) throughout the mixed-
liquor .
31
-------
TABLE 3
AVERAGE OPERATING CONDITIONS
PHASE I OPERATION (8/30/70-10/25/70)
FEED TO PLANT
Raw Wastewater Feed Rate (MGD)
Average for Period 2.79
Maximum Weekly Average 3.18
Dry Weather Daily Peak 3.5-4.0
Storm Flow Peak ~7.0
Raw Wastewater Temperature (°F) 68
Raw Wastewater pH 7-3
Raw Wastewater Solids Analysis (mg/1)
TSS 153
TDS 612
VSS 118
VDS 152
VSS/TSS 0.77
Raw Wastewater Substrate Analysis (mg/1)
BOD 218
COD 397
Raw Wastewater Nutrient Analysis (mg/1)
Ortho-Phosphorus as P 8.9
Total Phosphorus as P 12.6
NH3-N as N 21.6
TKN as N 34.9
N02-N as N 0.4
N03-N as N 0.8
OXYGENATION TANKS
Nominal Aeration Detention Time (hr) (Wastewater Only) 1.40
Aeration Detention Time (hr) (Wastewater + Recycle) 0.98
% Sludge Recycle 43
MLSS (mg/1) 5080
MLVSS (mg/1) 3680
Mixed-Liquor D.O. (mg/1) g.2
Food/Biomass Loading (Ib BOD applied/day/lb MLVSS) 1.01
Volumetric Organic Loading (Ib BOD applied/day/1000 ft3
mixed-liquor) 231
Overall 7, of Feed Oxygen Utilized 93 ^
32
-------
TABLE 3 CONTD1.
Based on Based on
Total Clarifier Upflow Annulus
Surface Area Surface Area
CLARIFICATION of 1260 ft2 of 820 ft2
Clarifier Overflow Rate (gpd/ft2)
Average for Period 1110 1700
Maximum Weekly Average 1260 1940
Dry Weather Daily Peak 1480 2280
Storm Flow Peak 2770 4280
Clarifier Liquid Rise Velocity (ft/hr)
Average for Period 6.1 9.5
Maximum Weekly Average 7.0 10.8
Dry Weather Daily Peak 8.2 12.7
Storm Flow Peak 15.5 24.0
Average Mass Loading (Ib TSS/day/ft2) 68.8
Sludge Recycle Rate (mgd) 1.2
Sludge Recycle Solids Analysis (mg/1)
TSS 21,600
VSS 14,900
33
-------
Mixed-liquor suspended solids were maintained at an average level of about
5100 mg/1 throughout the period with a range of weekly averages from 4600
to 5800 mg/1. As a consequence of maintaining this level of mixed-liquor
suspended solids, an average operating food-to-biotnass loading of about
1.0 pound of BOD applied per day per pound of MLVSS was achieved during
this phase of operation. The range of weekly average food-to-biomass
loadings varied from a low of 0.74 to a maximum of 1.23 pounds of BOD
applied/day/pound of MLVSS. The corresponding weekly average volumetric
loadings ranged from 181 to 271 pounds of BOD applied/day/1000 ft of
aeration tank volume with an average value of 231.
During Phase I, the system attained an average oxygen utilization of
93.47., with a minimum weekly utilization of 92.2%. The effluent vent
gas was regularly monitored at a composition of about 55% oxygen.
Both clarifiers at Batavia were used during this phase of operation.
The mixed-liquor flow from the oxygenation bay was split equally to
each clarifier. Underflow solids from both clarifiers were mixed and
returned to the inlet end of the oxygenation bay. Solids were wasted
daily from the clarifiers. The amount of wasting was determined by
estimating the plant influent load and by monitoring the position of
the sludge blanket in each clarifier. The clarifiers were operated to
hold the sludge blanket level between 10 and 12 feet from the top of
the clarifier during the early morning period of each day. This method
of control resulted in a low solids inventory in the clarifier and pro- ,
vided some storage capacity to handle the daily diurnal peak loads which
occurred in the day-time hours.
The Batavia clarifiers are designed as center feed, peripheral withdrawal
units. The center well of these clarifiers occupies 35 percent of the
total surface area of the clarifiers. The skirt which forms the center
well extends 7 feet-10 inches down into the liquid. Consequently, these
clarifiers operate as downflow clarifiers on the inner side of the center
wells and upflow clarifiers in the annulus between the center wells and
the takeoff weirs. Further details on the design of these clarifiers are
included in Appendix A of this report.
During Phase I, the average flow to the Batavia plant was 2.79 MGD.
This resulted in an average overflow rate through each clarifier of
1110 gpd/ft2 based on the total surface area of the clarifiers. The
equivalent rise velocities in each clarifier were 6.1 ft/hr based on
total surface area and 9.5 ft/hr based on the annular surface area.
The normal diurnal variations at Batavia resulted in peak dry weather
flows of 3.5 to 4.0 MGD during the day-time hours of moat weekdays.
Under these flow conditions, clarifier overflow rates and rise velocities
were higher than indicated above for the average flow conditions. Based
on an average diurnal peak of 35 percent above the average flow, an over-
flow rate of 1480 gpd/ft2 and a rise velocity of 8.2 ft/hr were regularly
experienced in the clarifiers based on total surface area. The rise
34
-------
velocity in the annular sections of the clarifiers during these periods
reached 12.7 ft/hr. Of course, substantially higher hydraulic loadings
were imposed during the infrequent storm flow periods when influent flows
as high as 7.0 MGD were experienced.
OPERATING CONDITIONS - PHASE II
During the last week of October, 1970, one-half of the total raw waste-
water flow to the Batavia plant was diverted into Bay 2 for treatment
using air aeration rather than oxygen aeration. The remaining one-half
of the total flow was directed through Bay 4 where oxygen aeration was
continued. The two clarifiers and sludge return systems were isolated
so that the air aeration and oxygenation systems were operating completely
independently. All recycle streams to the bays were split and the actual
compositions and flows to the oxygenation bay were carefully monitored
so that additional mass balances and sludge production data could be i
determined. Figure 4 shows a flow schematic of Phase II operation.
The operation of the plant during November was characterized by abnormally
high flows and correspondingly weak waste concentrations. Rainfall in
the area was high throughout the month and since infiltration at Batavia
is a regular occurrence, the situation persisted during this entire
period of operation. The average flow through the oxygenation bay, and
the single clarifier handling this bay, was 2.14 MGD for the four-week
operating period.
At the average flow rate, the resultant detention time in the oxygenation
bay was 1.8 hours based on raw wastewater flow. Table 4 presents a
summary of the wastewater feed characteristics as well as a description
of the operating conditions of the plant during this period. Additional
details relating weekly average values for all process operating parameters
are contained in Appendix D of this report.
During Phase II, the mixed-liquor suspended solids were maintained at an
average concentration of 3650 mg/1, and sludge was recycled at an average
rate of 36 percent of the raw wastewater feed rate. This inventory of
solids under aeration resulted in an average food-to-biomass loading of
0.63 pound BOD applied/day/pound MLVSS. This represented a food-to-
biomass loading intermediate to the range of loadings maintained in the
initial contract and, consequently, provided valuable additional sludge
production data.
The hydraulic load on the clarifiers was high during the entire Phase II
operation due to the persistent storm flow conditions which prevailed
during November. The average oxygenation clarifier overflow rate for
the period was 1700 gpd/ft based on the total clarifier surface area.
At this average flow condition, the equivalent rise rate in the clarifier
was 9.5 ft/hr based on the total surface area, or 14.4 ft/hr for the
35
-------
4.28 MGD
RAW WASTEWATER
(AVERAGE FOR PHASE H )
OXYGEN SYSTEM
>.I4 MGD j
1
1
1
BAY
r
1
•
4
^
NOT
IN
SERVICE
BAY 3
— -\
AIR SYSTEM
* 2.14 MGD
LJ
O
O
ID
_J
CO
CL
13
UJ
cr
BAY
LJ
o
Q
=1
'BAY
AIR
SYSTEM
-O
CLARIFIER
EFFLUENT
EFFLUENT
FIGURE 4
BATAVIA TREATMENT SYSTEM
PHASE H OPERATION
1970 CONTRACT
-------
TABLE 4
AVERAGE OPERATING CONDITIONS
PHASE II OPERATION (11/1/70- 11/30/70)
FEED TO PLANT
Raw Wastewater Feed Rate (MGD)
Average for Period ?.14
Maximum Weekly Average 2.22
Dry Weather Daily Peak
Storm Flow Peak ~7.0
Raw Wa^tewater Temperature (°F) 62
Raw Wastewater pH 7.1
Raw Wastewater Solids Analysis (mg/1)
TSS 83
TDS 661
VSS 63
VDS 103
VSS/TSS 0,76
Raw Wastewater Substrate Analysis (mg/1)
BOD 122
COD 177
Raw Wastewater Nutrient Analysis (mg/1)
Ortho-Phosphorus as P 3.7
Total Phosphorus as P 6.9
NH3-N as N 10.0
TKN as N 18.7
N02-N as N 0.5
N03-N as N 2.7
OXYGENATION TANKS
Nominal Aeration Detention Time (hr) (Wastewater Only) 1.82
Aeration Detention Time (hr) (Wastewater + Recycle) 1.34
70 Sludge Recycle 36
MLSS (mg/1) 3650
MLVSS (mg/1) 2580
Mixed-Liquor D.O. (mg/1) 12.1
Food/Biomass Loading (Ib BOD applied/day/lb MLVSS) 0.63
Volumetric Organic Loading (Ib BOD applied/day/1000 ft3
mixed-liquor) 101
Overall 70 of Feed Oxygen Utilized 91.3
37
-------
TABLE 4 CONTD.
CLARIFICATION
Clarifier Overflow Rate (gpd/ft2)
Average for Period
Maximum Weekly Average
Dry Weather Daily Peak
Storm Flow Peak
Based On
Total Clarifier
Surface Area
of 1260 ft2
1700
1760
2770
Based On
Upflow Annulus
Surface Area
of 820 ft2
2600
2700
4280
Clarifier Liquid Rise Velocity (ft/hr)
Average for Period 9.5
Maximum Weekly Average 9.8
Dry Weather Daily Peak
Storm Flow Peak 15.5
14.4
15.0
24.0
Average Mass Loading (Ib TSS/day/ft2) 73.0
Sludge Recycle Rate (mgd) 0.77
Sludge Recycle Solids Analysis (mg/1)
TSS 21,100
VSS 13,300
38
-------
annular section only. The clarifier is designed for a nominal flow
of 1.25 MGD. However, during Phase II, the flow to the clarifier
exceeded this rate 100 percent of the time and, in fact, exceeded
the total plant design flow of 2.5 MGD thirty percent of the time.
Waste solids at Batavia are disposed of by anaerobic digestion followed
by vacuum filtration and land filling. During the initial contract work
and during Phase I of this contract, this procedure was routinely followed
by the Batavia operators. During Phase II of this contract, however, no
vacuum filtration operations were conducted. Consequently, the amount of
supernatant recycled from the anaerobic digesters to the aeration bays
was increased. The effect of vacuum filtration is clearly evident in the
mass balances determined for the plant and in the comparative nutrient
removal levels achieved during each phase of operation. Both of these
topics will be discussed in the following sections of this report.
MASS BALANCES
A solids balance is fundamentally an imperfect basis for evaluating the
accuracy of the measurement of the performance parameters of a biological
waste treatment system. A portion of the gaseous oxygen entering the
system is incorporated into the cell mass formed and thereby converted
to solid form. Part of the organic carbon contained in the raw waste-
water is oxidized to carbon dioxide during treatment. The carbon dioxide
thus formed leaves the system in both the gaseous and dissolved forms.
Part of the hydrogen contained in the organic material in the raw waste-
water is oxidized to water during treatment. Because of the inter-
conversion of gases, liquids, and solids, no simple accounting of the
solids entering and leaving a biological treatment system can theoretically
balance. Therefore, a mass balance, as well as a solids balance, was made
to evaluate the accuracy of the measurement of performance parameters for
the oxygenation system. A mass balance differs from a solids balance in
that it also accounts for oxygen, carbon, and hydrogen entering or leaving
the system in gaseous or liquid form. Reference is made to the discussion
in the evaluation methods section of this report for all details on how
the solids and mass balances were made on the system, including the
analytical and sampling procedures as well as the assumptions. A short
summary of the approach is given below.
By referring to Figure 2, it can be seen that a simple accounting of the
raw wastewater entering the plant, the clarifier overflow stream, and
the waste sludge from the clarifier is not sufficient to fully account
for mass profiles through the system. In the normal operation of the
Batavia plant, liquid streams are recycled from other unit operations
at the site. These recycle streams include:
1. Return flow from the chlorine contact tanks which are
periodically cleaned by flushing the settled solids back
to the aeration bays.
39
-------
2. Thickener supernatant, including waste activated sludge
decant plus a continuous small flow of clarifier effluent
utilized as thickener wash water, which is recycled back
to the aeration tanks.
3, Vacuum filtrate which is recycled back to the aeration bays.
4. Digester supernatant which is transferred daily from the
secondary anaerobic digester back to the aeration bays.
All of the above streams contribute an additional solids, substrate, and
nutrient load to the aeration tanks and clarifiers which must be account
for. Since a prime objective of the present contract was to account for
solids profiles in the system, detailed flow and solids analyses were
performed on each of these recycle streams throughout the contract. This
was accomplished by measuring as well as possible the total flow of each
stream during each period when recycling was occurring. In addition,
composite grab samples were taken and analyzed for solids contents. The
results obtained were used for preparation of the mass and solids balances.
An accounting was made of the gas streams entering and leaving the process
by monitoring the influent flow of oxygen and the flow and oxygen compo-
sition of the effluent gas stream. Soluble gases were accounted for by
regular measurements of the effluent liquid pH as well as by periodic
determination of the wastewater alkalinity.
A summary of the weekly solids and mass balances for the oxygenation
system are presented in Table 5. This table shows the ratio of solids
and masses leaving the system to the same parameters entering the
system. The detailed weekly solids and mass balances on which Table 5
is based are presented in Appendix E of this report.
Table 5 indicates that the mass balance work was very successful in
accounting for mass entering and leaving the system. During both
periods of operation, a mass balance was closed to about 98 to 99
percent. An analysis of the solids profiles as summarized in Table
5 indicates the following:
1. The ratio of inert solids leaving to inert solids entering is
above 1.0 during Phase I and below 1.0 during Phase II. How-
ever, the deviation from 1.0 is not very great, indicating that
a good accounting of inert solids was obtained.
2. The Phase 1 data indicate that a portion of the inert suspended
solids entering in the feed wastewater are solubilized and
appear in the plant effluent and waste sludge effluent streams
in soluble form. During Phase II, the ratios of ISS and IDS
were both below 1.0 which does not support the solubilization
theory.
40
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TABLE 5
SUMMARY OF SOLIDS AND MASS BALANCES
II
FOR BATAVIA OXYGENATION SYSTEM
Ratio of Quantities Leaving to Quantities Entering
WEEK OF
9-6-70
9-13-70
9-20-70
9-27-70
10-4-70
10-11-70
10-18-70
Average
11-1-70
11-8-70
11-15-70
Average
TSS VSS
0.77 0.76
1.03 0.90
0.91 0.87
0.95 1.11
1.15 1.29
0.93 0.98
1.00 1.15
0.95 0.97
0.60 0.54
0.97 1.05
1.13 1.08
0.89 0.88
LEGEND
TSS
VSS
ISS
TDS
VDS
IDS
TS
VS
IS
Mass
ISS TDS VDS IDS TS VS IS
.82 0.99 0.76 1.06 0.93 0.76 1.04
1.34 0.98 0.67 1.10 1.00 0.80 1.14
1.03 1.00 0.49 1.17 0.98 0.67 1.16
0.68 1.07 0.91 1.11 1.04 1.00 1.06
0.81 1.00 0.70 1.11 1.02 0.90 1.09
0.80 1.00 0.60 1.11 0.98 0.78 1.08
0.70 0.98 0.70 1.06 0.98 0.90 1.02
0.89 1.00 0.69 1.10 0.99 0.82 1.08
0.79 0.97 1.01 0.96 0.90 0.76 0.95
0.78 0.90 0.58 1.00 0.91 0.79 0.98
1.25 0.97 1.16 0.95 1.00 1.12 0.97
0.92 0.95 0.85 0.97 0.94 0.87 0.96
= Total suspended solids
= Volatile suspended solids
= Inert suspended solid.s
= Total dissolved solids
= Volatile dissolved solids
= Inert dissolved solids
= Total solids
= Volatile solids
Inert solids
= Total solids + delta C + delta 0, + delta 11,
MASS
0.93
1.15
L.02
0.96
0.92
0.97
0.94
0.98
0.93
1.04
1.00
0.99
41
-------
3. The volatile solids decreased between the influent and
effluent. This is expected as substrate is metabolized
by the biomass to produce carbon dioxide and water.
The detailed weekly mass and solid balances can be studied by referring
to Appendix E. Table 6 contains an average of these detailed weekly
balances for each phase of operation. From this table it can be seen
that:
1. The recycle streams do contribute a significant fraction of
mass to the total oxygenation bay inlet mass. The major
contribution is from the digester supernatant in the form
of suspended solids.
2. A major portion of the wastewater influent solids occurs as
dissolved solids and a large fraction of these solids are
inert.
3. Sludge wasting does not contribute substantially to the
removal of inert solids since the majority of inert solids
enter and leave as soluble material.
4. The absence of vacuum filtering during the Phase II operation
in November resulted in a significant increase in the digester
supernatant solids recycled back to the oxygenation bay. The
average total solids transferred back during September and
October were 822 pounds per day, while during November the
amount transferred back increased to 1045 pounds per day.
The overall effect on the Batavia operation is especially
significant since the November transfer figure represents only
one half of the total. The other half was transferred to the
bays treating fifty percent of the wastewater with air.
42
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TABLE 6
AVERAGE WEEKLY BATAVIA MASS BALANCES FOR PHASES I AND II
TSS
VSS
TDS
VDS
IS
VS
IS
AC
A02
AH,
Mass
TSS
VSS
TDS
VDS
TS
VS
IS
AC
AO,
AM2
Mass
Feed Raw
Oxygen Wastewater
3350
2572
13,518
3307
16,867
5879
10,988
5098
5098 16,867
Feed Raw
Oxygen Wastewater
1513
1240
11,462
1913
12,978
3153
9825
3480
3480 12,978
Mass In - TS
Founds Per Day In
Digester Vacuum
Supernatant Filtrate
537 157
330 62
50 78
23 16
587 235
353 77
234 158
587 235
Pounds Per Day In
Digester Vacuum
Supernatant Filtrate
979
622
67
38
1045
660
386
1045 0
in + A0'ln
PHASE I
C12 Tank
Transfer
120
21
36
6
156
27
128
156
PHASE II
Clj Tank
Transfer
60
10
18
3
78
13
65
78
, Pounds Per Day Out
Total
Ib/day In
4164
2985
13,682
3352
17,845
6336
11,508
5098
22,943
Waste Clarlfler
Oxygen Effluent
652
527
13,577
2274
14,229
2800
11,429
19 989
381 3307
133
400 18,658
Waste
Sludge
3306
2381
134
29
3441
2410
1030
3441
Total
Ib/day Out
3958
2908
13,711
2303
17,670
5210
12,459
1008
3688
133
22,499
(Jill
In
0.95
0.97
1.00
0.69
0.99
0.82
1.08
0. 72
II. 'IH
Pounds Per Day Out
Total
Ib/day In
2552
1872
11,547
1954
14,101
3826
10,276
3480
17,581
MdNH
Waste Clarlfler
Oxygen Effluent
829
662
10,912
1657
11,741
2320
9404
29 839
410 2807
115
439 1.5,502
Out - TS + AC
ouL ui
Waste
Sludge
1446
986
48
10
1494
993
501
1494
Total
Ib/day Out
2275
1648
10,960
1667
13,235
3313
9905
868
3217
1 15
I7,43r>
Out
In
0.89
0.88
0.95
O.B5
0.94
0.87
0.96
0.92
il.'l'l
, + AO, >• .Ml,
n JouL 'mil
43
-------
EXCESS SLUDGE PRODUCTION
The solids balances developed for each week of the contract were used
to determine the excess sludge production for the comparable periods
during the contract. These excess sludge production figures are
considered to be very representative based on the close mass balances
achieved in both Phase I and Phase II. The sludge production figures
summarized below were obtained directly from the weekly solids balances
in Appendix E.
Week of
PHASE I
9- 6-70
9-13-70
9-20-70
9-27-70
10- 4-70
10-11-70
10-18-70
PHASE II
11- 1-70
11- 8-70
11-15-70
VSS in Clarifier
Effluent
(Ib/day)
845
378
241
481
916
443
383
398
707
982
VSS in
Waste Sludge
(Ib/day)
2280
3141
2287
2179
1754
2325
2699
745
1433
781
Total Excess VSS
(Ib/day)
3125
3519
2528
2660
2670
2768
3082
1043
2140
1763
The excess sludge production data from the Phase I operation were ob-
tained under similar conditions to those prevailing in the initial
contract work at Batavia and, therefore, can be compared directly with
that previous data. The Phase II results, however, cannot be directly
compared with the 1969 or the Phase 1-1970 work since the vacuum filter
was not operated during the period. As shown earlier, a significant
increase in recycle solids from the digester occurred during November
as a consequence of not removing solids from the treatment plant. The
increased recycle of supernatant from the digester was also responsible
for a significant decrease in net removal of nutrients by the process.
In order to make any comparison of excess solids production during
November with previous work, the following adjustments were necessary
to the November excess sludge production figures.
1. The solids returned to the aeration bays from the anaerobic
digester must be considered to be quite highly stabilized„
44
-------
2.
Thus, they would be expected to pass through the system
without substantial biological degradation. For this
analysis, it has been assumed that 30 percent of the
volatile solids in the digester supernatant would be
degraded.
The incremental solids recycled from the digester in
November relative to September/October would not have
occurred had the filtering operation been in effect.
These incremental solids, therefore, were discounted from
the total excess sludge production figures for November as
follows:
Week
of
Super-
natant
VSS
Recycled
(Ib/day)
Average
VSS
Recycled
Sept. -Oct.
(Ib/day)
November
Increment
(Ib/day)
70% of
November
Increment
(Ib/day)
Total
November
Excess
VSS Formed
(Ib/day)
Adjusted
November
Excess VSS
(Ib/day)
11- 1
11- 8
11-15
517
667
681
330
330
330
187
337
351
130
235
245
1043
2140
1763
913
1905
1518
With the above adjustment, it was possible to determine the excess sludge
produced per day per unit of suspended solids inventory by the oxygena-
tion system, for both phases of operation. The BOD removed per day per
unit of volatile suspended solids inventory could also be determined.
These figures are tabulated as follows:
Week of
PHASE I
9- 6-70
9-13-70
9-20-70
9-27-70
10- 4-70
10-11-70
10-18-70
PHASE II
11- 1-70
11- 8-70
11-15-70
Ib VSS Formed/day
Ib MLVSS
Under Oxygenation
0.58
0.74
0.46
0.57
0.50
0.57
0.57
0.24
0.52
0.43
Ib BOD Removed/day
Ib MLVSS
Under Oxygenation
0.93
1.17
0.68
0.79
0.84
1.02
1.05
0.40
0.62
0.65
45
-------
It was possible to compare these sludge production figures with data
obtained during the initial contract. Excess sludge production was
correlated in the initial contract report by independently plotting
excess volatile solids formation and BOD removal data as a function
of aeration detention times for both the air aeration and oxygenation
systems. These two curves are reproduced in this report as Figures 5
and 6. They were then used to generate a sludge production relationship
commonly employed in biokinetic design and evaluation work. This rela-
tionship is presumed to take the form of a straight line with a slope
commonly referred to as the yield coefficient and an intercept represent-
ing a cell decay or endogenous respiration coefficient. This curve is
also reproduced as Figure 7.
This indirect approach to correlating the data was taken in the initial
contract because attempts at correlating the data directly did not yield
useful correlations. It is apparent from observation of Figures 5 and
6 that substantial scatter does exist. During the present work, additional
sludge production data was obtained for the oxygenation system as has been
discussed above. This additional data was obtained at conditions which
allow extension of the data range for correlation. On the basis of the
data obtained in 1970, a revised correlation was made using both the 1969
and 1970 data. It was found that a direct correlation could now be made
although scatter is still evident. The results of this correlation are
shown in Figure 8.
Figure 8 indicates that the 1970 sludge production data confirms the
sludge production data obtained in 1969 except that the slope of the
oxygen correlation is not as steep as indicated in the original report.
The air system correlation is substantially the same except for a slight
decrease in the slope of the curve. Examination of the two curves indic-
ates that an oxygenation system can be expected to produce substantially
less solids than would be expected from a conventional air aeration system
throughout the normal design range of an activated sludge system. It is
evident, however, that the relative magnitude of the decreased solids
production becomes less as the specific reaction rate becomes smaller.
Since the 1970 sludge production data permitted a revision of the sludge
production correlation, the cost estimates developed under the original
contract have also been revised. The revised cost analysis is discussed
later in this report.
MISCELLANEOUS RESULTS
The preceding portion of this discussion has concentrated on two of the
major contract objectives - further elucidation of solids profiles and
sludge production characteristics of an oxygenation system. In the
course of operating the oxygenation system to obtain the data necessary
to meet those objectives, a large quantity of information on substrate
and nutrient profiles as well as other general operating characteristics
of an oxygenation system were obtained. Since this information was
46
-------
1.2-
to
CD
1
O
UJ
2
cr
o
CO
o
X
UJ
m
1.0
0.8
0.6
0.4
0.2-
AIR SYSTEM
PHASE I OPERATION A
PHASE m OPERATION O
~WYGEN SYSTEM
PHASE I OPERATION
PHASE E OPERATION
O
u 12345
AERATION DETENTION TIME (HR)
FIGURE 5
COMPARATIVE CORRELATION OF EXCESS VSS
AND MLVSS WITH AERATION DETENTION TIME
FROM INITIAL CONTRACT REPORT 17050 DNW 05/70
47
-------
in
03
o
UJ
cc
o
o
CD
CD
1.0
0.8
0.6
0.4
0.2
AIR SYSTEM
PHASE I D
PHASE HI •
OXYGEN SYSTEM
PHASE I A
PHASE H O
PHASE IT •
234
AERATION DETENTION TIME (HR)
OXYGEN
FIGURE 6
COMPARATIVE CORRELATION OF BOD REMOVED
AND MLVSS WITH AERATION DETENTION TIME
FROM INITIAL CONTRACT REPORT 17050 DNW 05/7Q
48
-------
CO
co
CD
o
UJ
s
tr
o
co
CO
en
CO
UJ
o
X
UJ
CD
1.0-
0.8n
0.6
0.4-
0.2-
AIR SYSTEM \j = I.38X-O.I7
OXYGEN SYSTEM U = I.05X-0.27
AIR
/ 0.2/
0.4
0.6
0.8
1.0
^r
-0.17-f / LB BOD REMOVED/DAY/LB MLVSS
-0.27-K
FIGURE 7
CORRELATION OF EXCESS VSS FROM
AIR AERATION AND OXYGENATION SYSTEMS WITH
BOD REMOVAL AND MLVSS
FROM INITIAL CONTRACT REPORT 17050 DWN 05/70
-------
v>
v>
o
1.0
0.9
0.8
0.7
0.6
1 0.5
w
w O.1
UJ
u
X
111
0.3
FIGURE 8
EXCESS VSS FORMATION
CORRELATIONS AND
95% CONFIDENCE LIMITS
FOR PREDICTED VALUES
1969 AND 1970 /
BATAVIA OPERATIONS /
/ AIR AERATION SYSTEM
/
/
/o
OXYGENATION
SYSTTM
0.2
0.
O
/ D
CODE
A
A
D
•
O
•
SYSTEM
AIR
AIR
OXYGEN
OXYGEN
OXYGEN
OXYGEN
YEAR
1969
1969
1969
069
1970
1970
PHASE
T
m
i
TL
I
TL
i . i i
0.8 1.0
LB BOD REMOVED/DAY/LB MLVSS
.2
1.4
50
-------
interpreted quite extensively as part of the initial contract objectives,
it will not be treated in depth here. A few comments, however, are per-
tinent.
Filtered samples of the various streams were not analyzed during the
initial contract work. Therefore, it was not possible to determine the
performance characteristics of the oxygenation system independently from
the performance of the clarifiers. "During the present contract, soluble
as well as total substrate analyses were performed so that the performance
of the oxygenation system and the clarifiers could be determined separately,
Clarifier effluent soluble BOD concentrations averaged 14.0 mg/1 during
Phase I and 8.5 mg/1 during Phase II. There concentrations are indicative
that the oxygenation process can be operated at high loadings and short
retention times and still assimilate the BOD in the incoming wastewater
stream. The same conclusion is evident from the soluble COD data ob-
tained. During Phase I, the effluent sloluble COD averaged 61 mg/1,
while during Phase II the comparable value was 36 mg/1.
The total BOD in the effluent was higher than the soluble BOD due to
suspended solids carried over the clarifier weirs. During both phases
of the contract, periods of high hydraulic loading on the clarifiers
occurred causing extensive solids carry over. Since flow weighted
composite samples were collected continuously, these periods of high ,
flow have important effects on the total effluent suspended solids
levels. Generally, during phase I operation, solids loss occurred
during some of the day-time hydraulic peaks when clarifier surface
overflow rates exceeded 1500 gpd/ft and rise velocities in the annular
sections of the clarifiers reached 10-12 ft/hr. During Phase I, the <
average suspended solids concentration in the effluent was 32 mg/1 and
effluent total BOD averaged 40 mg/1. Throughout the entire period of
Phase II, the hydraulic load on the clarifiers was in excess of the
design flow so that periods did occur when solids were lost over the
weir in appreciable quantities. The average overflow rate for the
period was 1700 gpd/ft^ and clarifier rise velocities in the annular
section averaged 14-15 ft/hr. As a consequence, the average Phase II
effluent suspended solids concentration was 81 mg/1 and the resultant
average effluent total BOD was 41 mg/1. Because of the hydraulically
overloaded clarifier conditions, however, these effluent suspended solids
and BOD levels are not representative of the effluent quality to be ex-
pected under these oxygenation system operating conditions with adequate
clarifier capacity.
As was discussed earlier, the;recycle rate of anaerobic digester super-
natant was increased during Phase II because the plant vacuum filter was
not operated. The increased return of this supernatant material was
readily evident in,the plant operating results. Among the specific
results observed are the following.
During Phase I when a normal sludge filtration schedule was maintained,
the VSS/TSS ratio in the effluent was 0.79. However, during the Phase II
51
-------
operation this rate decreased to 0.69, even though the volatile fraction
in the feed remained nearly constant during both phases of operation
(0.75-0.77).
The oxygen consumption per pound of BOD removed increased in Phase II
compared with Phase I. The average value of 1.01 obtained during
Phase I confirmed the data determined in the initial contract work when
plant vacuum filtering operations were regularly conducted. The higher
value of 1.50 obtained in Phase II is a result of the increased load of
BOD returned to the system as anaerobic digester supernatant and not
accounted for as feed BOD.
During Phase I, total nitrogen removal by the system averaged 21.6
percent. This value was generally consistent with previous data ob-
tained in the initial contract. During Phase II, however, the average
removal of nitrogen actually decreased and became negative (-4.1 per-
cent) due to the high quantity of nitrogen contained in the digester
supernatant recycle to the oxygenation bay. During both phases of
operation, organic nitrogen in the feed was converted to ammonia
nitrogen as expected. Some traces of nitrite and nitrate nitrogen were
noted in the effluent; however, significant nitrification was not ob-
served, nor was it expected based on the high food-to-biomass loadings
maintained.
Both total and ortho-phosphorus data were taken during this contract.
Total phosphorus removal averaged 14.3 percent during Phase I. Ortho-
phosphorus comprised 70 to 80 percent of the total phosphorus in the
wastewater feed stream and 95 to 100 percent of the total phosphorous
in the effluent stream. The phosphorous profile through the plant
during Phase II also reflects the results of not operating the plant
vacuum filter during the period as the average total phosphorous re-
moval was zero.
ECONOMIC COMPARISON OF OXYGENATION AND CONVENTIONAL DIFFUSED AIR AERA-
TION TREATMENT
Figures 26 through 29 in the initial contract report compared total
wastewater treatment cost estimates using conventional diffused air
aeration and oxygenation systems at' various detention times and air
requirements. These estimates have been revised and are presented as
Figures 9 through 12 in this report. Such revision is necessary as a
result of the new sludge production estimates which are based on the
correlations in Figure 8 of this report. A second and equally import-
ant reason is the availability of more current cost information on
oxygenation systems after an additional year of experience in designing
and estimating such systems. All other factors and parameters includ-
ing costs used in drawing the comparison of conventional air aeration
versus oxygenation system economics are identical to those in the
initial report, and may be extracted from that report„
52
-------
Table 7 displays sludge production estimates for a 100 MGD plant for
both conventional air aeration and oxygenation systems, as predicted
by the correlations illustrated in Figure 8. These estimates indicate
that an oxygenation system produces from 38-44 percent less waste
activated sludge solids than a conventional air aeration system over
the range of detention times considered. This decreased production
of waste activated sludge is an important factor which reduces the
cost of an oxygenation system in comparison to a conventional air
aeration system. Another factor considered in this and the initial
report is the increased solids concentration in the clarifier under-
flow, which magnifies the difference in volume of waste activated sludge
between the two systems, and has significant impact on the economics of
the sludge dewatering equipment which must be included in either facility.
The dewatering characteristics of sludge from an oxygenation system, as
demonstrated by vacuum filtration tests reported herein, will also lead
to further economies.
As mentioned earlier, the other major factor which necessitated revision
of the economic analysis is the current cost of oxygenation systems.
Since issuance of the initial contract report, a good deal of activity
with respect to the marketing of oxygenation systems has taken place,
and a number of facilities incorporating oxygenation systems are in
advanced stages of design and construction. The increased level of
design activity has led to optimization of components used in the
oxygenation system, with a resultant decrease in costs. These current
costs are reflected in Figures 9 through 12 and are a natural consequence
of a maturing design. One facet of this optimization is the use of
surface aerators for oxygen transfer in most systems up to 30 MGD
employing relatively shallow tanks (up to 17 ft liquid depth). Plants
in this size range are relatively capital cost sensitive, and the use of
this equipment can sometimes lead to significant economies. Larger
plants employing deeper tanks are designed using the more efficient
sparger-impeller combination, which is particularly cost effective in
the larger deep tank systems and also affords greater power savings and a
reduced number of mechanical units.
The general conclusion which can be drawn from a study of Figures 9
through 12 is that the economic attractiveness relating to the use of
oxygen aeration in the activated sludge process has been further sub-
stantiated by the work documented in this report. A change from the
earlier economic comparison, however, is that the oxygenation system
treatment costs no longer decrease with increasing detention time,,
The updated correlation shows the magnitude of reduction in solids pro-
duction in the low food-to-biomass range is not as great as predicted by
the earlier correlations. Therefore, the figures presented in tliis
report show that oxygenation system treatment costs over the range of
commonly accepted detention times are relatively constant and actually
increase slightly with an increase in detention time in the oxygena-
tion tanks.
53
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TABLE 7
REVISED SUMMARY OF SLUDGE DISPOSAL ESTIMATES FOR
Ul
AIR
Aeration Detention
Time (Hr) MLVSS Cone.
(Based on Raw Flow) (mg/1)
3.0
4.0
5.0
6.0
1.00
1.33
1.66
2.00
1625
1625
1625
1625
4420
4420
4420
4420
AERATION AND OXYGEN AT I ON SYSTEM ECONOMIC COMPARI!
100 MGD PLANT
Total Plant Primary Sludge
Lb MLVSS Lb BOD Produced
Under Aeration Removed/Day (lb/day)
169,000
226,000
282,000
339,000
154,000
204,000
255,000
30P,000
AIR AERATION SYSTEM
150,000
153,000
156,000
158,000
OXYGENATION SYSTEM
150,000
153,000
156,000
158,000
86,000
86,000
86,000
86,000
86,000
86,000
86,000
86,000
*Total
Waste Activated
Sludge Produced
(lb/day)
150,000
150,000
149,000
147,000
84,000
87,000
89,000
91,000
Total Sludge
Produced
(lb/day)
236,000
236,000
235,000
233,000
170,000
173,000
175,000
177,000
Calculated from 1970 correlations.
-------
36
32
28
FIGURE 9
TOTAL TREATMENT COSTS
I MGD TREATED
CO
z
o
to
o
o
o
24
20
^ 16
co
fc
O
o
»-
111
<
Id
12
WITH AIR AERATION
CF (NTP) OF AIR PER
GALLON (PERLB BOD)
2.4 (2250)
1.6 (1500)
0.8 (750)
WITH OXYGENATION
J*
p
HOURS DETENTION IN OXYGENATORS
1.0 1.5 2.0
3456
HOURS DETENTION IN AIR AERATORS
55
-------
36
32
28
24
(9
O
O
O
20
' 16
O
O
LJ 12
FIGURE 10
TOTAL TREATMENT COSTS
6 MOD TREATED
WITH AIR AERATION
CF (NTP) OF AIR PER
GALLON (PER LBBOD)
2.4 (2250)
1.6 (1500)
0.8 (750)
WITH OXYGENATION
UJ
oc
HOURS DETENTION IN OXYGENATORS
1.0 1.5 2.0
3456
HOURS DETENTION IN AIR AERATORS
56
-------
36
32
28
FIGURE II
TOTAL TREATMENT COSTS
30 MGD TREATED
CO
3
24
o
O 20
O
o
I
(O
fc
o
o
(U
1
U
<
o
16
12
8
WITH AIR AERATION
CF (NTP)OF AIR PER
GALLON (PERLB BOD)
2.4 (2250)
1.6 ( 1500)
0.8 ( 750)
WITH OXYGENATION
HOURS DETENTION IN OXYGENATORS
1.0 1.5 2.0
3456
HOURS DETENTION IN AIR AERATORS
57
-------
16
16
FIGURE 12
TOTAL TREATMENT COSTS
100 MGD TREATED
14
w 12
O
(9
o io
o
o
o
ui
tr
6
WITH AIR AERATION
CF(NTP)AIR PER
GALLON(PER LB BOO)
2.4 (2250)
1.6 (1500)
0.8 (750)
WITH OXYGENATION
HOURS DETENTION IN OXYGENATORS
1.0 L5 2.0
3456
HOURS DETENTION IN AIR AERATORS
58
-------
A number of additional comments could be made on these figures in an
attempt to compare treatment costs using conventional diffused air
aeration and oxygenation systems. It is worthwhile to note, however,
that these curves are quite general, and a wide variation can exist
in the individual parameters, cost factors, and design conditions
used to construct them. As a result, the comparisons of such systems
may vary greatly from location to location, depending on the specific
design conditions and cost factors in question. As a general conclu-
sion, however, the overall economic attractiveness relating to the use
of oxygen aeration in the activated sludge process is amply demonstrated
by the data contained in this report, which combined with the initial
contract report resulted in the comparisons shown in Figures 9 through
12.
59
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SECTION VII
VACUUM FILTRATION
Results from the initial contract work at Batavia showed that, under
conditions of proper clarifier operation, oxygenation system waste
activated sludge concentrations could be maintained routinely at 2-3
percent. In a few laboratory tests, it was shown that the dewatering
characteristics of the oxygenated waste sludge were superior to that
wasted from a conventional ait aeration system. It was speculated that
it might be possible, then, to directly filter the oxygenation system
waste activated sludge without thickening and with a smaller chemical
requirement.
To substantiate these earlier observations, and to provide meaningful
data on the dewatering characteristics of oxygenation system waste
activated sludge, one of the objectives of this contract was to evalu-
ate the filterability of this sludge by periodically conducting direct
filtering tests using a pilot scale (3 ft diameter x 1 ft length) Eimco
Corporation rotary belt vacuum filter. The results of these vacuum
filtration studies are reported in this section.
A flow sheet of the experimental operation is given in Figure 13. After
preliminary testing, operation of this unit proceeded smoothly. It
usually took 2-3 hours for a complete run, including time for prepara-
tion of chemical conditioning solutions, adjustment of flow rates, and
data collection. In contrast, many plant-scale vacuum filter operations
are continuous and semi-automatic.
Additional tests were conducted using a standard filter leaf test
apparatus . This apparatus was used to ascertain the suitability of
various polymeric and inorganic chemicals as conditioners for use in
the pilot-scale vacuum filter. These tests are summarized in Table 3
of Appendix F.
The filter medium selected for both the Eimcobelt vacuum filter and the
filter leaf was a synthetic (o'rlon) cloth, Eimco OR 593 , based on
preliminary tests conducted on the filter leaf test apparatus. This
cloth was chosen both on the basis of cake release and filtrate quality.
In reviewing the sludge conditioning reports of 18 treatment plants",
it was found that the use of ferric chloride and lime constitutes norm;i I
practice for plants dewatering raw and digested sludges, while ferric
chloride alone is normally used for those plants dewatering waste
activated sludge.
Vacuum filtration studies were conducted using the Eimcobelt filter on
anaerobically digested sludge from the Batavia digesters, on the
61
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FIGURE 13
FLOW SHEET OF VACUUM FILTER APPARATUS
SLUDGE
PUMP
-RAW WASTE ACTIVATED
-ANAEROBICALLY DIGESTED
WASTE ACTIVATED
-AEROBICALLY DIGESTED
WASTE ACTIVATED
DEWATERED
SLUDGE
FILTRATE
FILTRATE PUMP
-------
oxygenation system waste activated sludge taken directly from the under-
flow of the clarifiers, on thickened waste activated sludge from the
oxygenation system, and finally on aerobically (oxygen aerated) digested
waste activated sludge from the oxygenation system. The results of
these studies are tabulated in Tables 1 to 5 in Appendix F. All the
runs were made under the following conditions:
Parameter Sludge Type
Anaerobic 02 Digested Waste Activated
Temperature (°C) 25-30 18-25 20-25
PH 6.5-7.5 6.5-7.0 6.0-7.0
Alkalinity (mg/1 as CaC03) - 1000-1500 650-750
Drum submergence = 25-30%
Filter Vacuum = 15-20 in. Hg
Actual Running Time = 20-60 min.
From these studies, the following conclusions were drawn:
1. Oxygenated waste activated sludge can be filtered directly without
thickening. It is believed that this is a unique feature of the
oxygenated sludge not generally achievable with activated sludge
from conventional air aeration systems. The cake yield obtained
on unthickened oxygenated waste activated sludge varied from 1.0
Ib/hr/ft2 with 1.0% TSS in feed to 4.5 Ib/hr/ft with 2.0 to 2.5%
TSS in feed at a cycle time of 2.5 minutes.
2. Oxygenated waste activated sludge can be thickened up to 4.5% by
simple gravity settling and then filtered to increase cake yield
to 5.1 Ib/hr/ft2.
3. For the oxygenated waste activated sludge at Batavia, the sludge
cake yield, L, can be correlated by the following equation:
fr 0.767
L = 3.85
0.656
where; C- = Initial total solids concentration in the feed sludge
(before conditioning) in %, solids by wt.
L = Sludge cake yield in pounds of dry total aolids in the
sludge cake discharged from the filter per hour per
square foot of total filter urea.
t = Total cycle time (time required for one complete
revolution of a filter drum) in minutes/revolution.
63
-------
From the above equation, it can be seen that the yield varies (to
the exponential powers shown) directly with initial solids concentra-
tion and inversely with cycle time.
4. The-most effective chemical conditioner for the waste activated
sludge was ferric chloride with the optimum dosage rate of 200 pounds
of ferric chloride per ton of dry solids, which is equivalent to an
approximate chemical cost of $8 per ton of dry solids filtered. At
this dosage rate, the pH of the conditioned sludge fell to around
3.0. Therefore, it may be possible to control the addition rate of
ferric chloride in a plant size operation by pH.
5. None of the polymers tested were as effective as ferric chloride
for conditioning oxygenated waste activated sludge for vacuum fil-
tration under the conditions tested 3t Batavia in this contract.
6. It is much more difficult to vacuum filter aerobically digested
sludge than waste activated sludge.
7. The combination of ferric chloride and lime was the most effective
conditioner for vacuum filtration of anaerobically digested sludge.
OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE
A number of sludge and operating variables affect vacuum filtration
performance'' . The sludge variables are solids concentration, sludge
and filtrate viscosity, sludge compressibility, chemical composition,
and the nature of the sludge. The operating variables are the filtration
vacuum, drum submergence, drum speed, filter media, and chemical
conditioners used prior to filtration.
The theory of filtration has not yet developed into an exact science due
to its complexity. Nevertheless, understanding of the filtration theory
greatly helps in the interpretation of experimental data and in the
prediction of the rate of cake formation as a function of various filtra-
tion parameters mentioned previously.
The operation of a continuous filtration device for the dewatering of
sludge is a cyclic process . It encompasses many steps or functions
such as the mixing of sludge with chemical conditioners, the formation
of sludge cake, the drying of the sludge cake formed, and finally cake
discharge from the filter. Any one of these steps might be rate
controlling; yet all must interact and, therefore, none can be ignored.
The dewatering or drying r.ite of a vacuum filter sludge cake has been
found to be a complex phenomenon which is less amendable to theoretical
treatment than the rate of cake formation^'^• The major factors known
to influence filter cake moisture content are**:
64
-------
1. Cycle time, or rather drying time.
2. Vacuum of drying zone.
3. Filtrate viscosity.
4. Feed solids concentration and chemical conditioning.
5. Type of sludge.
6. Cake thickness (which might be a dependent variable) .
The variable that has the greatest effect on cake moisture for a given
sludge and which can commonly be controlled in a filter operation is
drying or cycle time. Figure 14 shows a plot of cake moisture (expressed
as Ib H20 per Ib of dry solids in the cake) versus cycle time, at con-
stant submergence, constant pressure, and relatively constant feed
solids concentration from data at Batavia with oxygenated waste activated
sludge. As is expected, the cake moisture decreases as cycle time in-
creases. The cake yield or filtration rate, on the other hand, also
decreases as the cycle time increases. There is obviously a trade-off
between the moisture content in the cake and the cake yield. Therefore,
after the cake has been sufficiently dewatered for the subsequent
disposal step(s), little if any benefit is gained by a longer cycle or
additional drying time.
The effect of feed solids concentration on cake moisture content is,
on the other hand, not as apparent. The data taken with oxygenated
waste activated sludge at Batavia failed to yield a meaningful correla-
tion between these two parameters.
A list of nomenclature used in the following development is contained
in Table 6 at the end of Appendix F. The liquid passing through the
filter cake has been assumed to follow streamline motion so that the
modified Poiseuill differential equation for the flow of liquid through
capillaries applies"' *-® . Equation 1 gives the instantaneous rate of
filtrate collected during the cake formation cycle of the filtration
operation:
Ade
•3
where: V = Volume of filtrated collected (ftj)
f\
A = Area of filtering surface (ff4)
9 = Form time (seconds or minutes)
P = Pressure drop through filter medium and sludge cake
(Ib force/ft^)
65
-------
10
9
8
7
UJ 6
o
(E J
0
00
_l
a 4
(VI
CO
FIGURE 14
OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE
CAKE MOISTURE VS. CYCLE TIME
BASIS
VACUUM. !5-20"Hg
SUBMERGENCE: 25-30%
% SOLIDS IN FEED: 2.0-2.4
FeCI3 ADDED AS % (BY WT) DRY TOTAL SOLIDS: 8-13
FILTRATION AREA: 9.0 FT2
= 7.74X~0367 (CORRELATION COEFFICIENT =0.868)
O
uj 3
o
z
UJ
1C
o
6 2
_u
2 345
CYCLE TIME (MINUTES/REVOLUTION)
8 9 10
-------
n
gc = Newton's law conversion factor (ft-lb/lb force-sec )
fJi = Viscosity of filtrate (Ib/ft-sec)
Q( = Average specific cake resistance (ft/lb)
w = Weight of dry cake per unit volume of filtrate (lb/ft3)
r = Resistance of filter medium (ft"1)
Since the pressure drop through a cake is proportional to the velocity
of flow (laminar flow), Equation' 1 means that the filtration rate is
equal to the driving force, i.e., the pressure, divided by the resistance
This resistance consists of two parts: the resistance of the cake it-
self (represented by a, the specific cake resistance), and that of the
filter medium, r. The specific resistance is a function of many factors,
but is usually a constant for one slurry, i.e., one sludge type condi-
tioned by the same chemical^>7,8,11. xhe resistance of the filter
medium, with proper vacuum filter design, can be considered negligible,
and will be treated as such. Therefore, Equation 1 can be integrated
at constant pressure:
de (2)
W.OCW
(3)
u « w
or /
Filtrate Rate = -I =/ 2Pgr \ 1/2 (4)
Ae m t* w 6
But W, the weight of dry cake solids formed during any time period, is
equal to wV so that:
Cake Formation Rate =' *_ = /WV/2 (5)
Ae lu^e
Let L = yield of cake (as dry solids)
S = 7o drum submergence
t = total cycle time
Therefore,
' _ tS_ (6)
H 100
67
-------
and L = — (7)
At
Substitute Equations 6 and 7 into Equation 5:
= 100 200gcPw
s I u * s t y v '
But w is proportional to the % solids (by wt . ) in the feed sludge, C^.
Therefore, for constant pressure and constant drum F .bmergence (as in our
studies), the filtration yield (of cake) is:
1/2
(9)
V *•/
where :
K = a constant
However, the exponents for cycle time and initial solids concentration
deviate from Equations 8 and 9 for most wastewater sludges6'8 > ^2 . It
is recommended*^ that Equation 9 be modified to:
r m
i } (10)
tn
where the exponents m and n are experimentally obtained.
According to Equation 10, at constant pressure and drum submergence, the
cake yield of a (conditioned) sludge is a function of the initial solids
concentration in the sludge before filtration and the cycle time of fil-
tration only.
Figure 15 shows the cake yield, L, plotted on logarithmic scales against
the percent (by weight) solids in the feed, C^, at a constant cycle time,
t, from data obtained with oxygenated waste activated sludge conditioned
with Fed-} only. The data obtained for vacuum filtration of oxygenated
waste activated sludge are contained in Tables 1 and 2 of Appendix F.
Table 1 contains the data for waste activated sludge filtered without
thickening, while Table 2 contains the data for runs with thickening.
The slope of the plot, m, = 0.767, with a standard deviation of 0.0411
and a correlation coefficient of 0.970. Similarly, Figure 16 shows the
cake yield, L, plotted against cycle time, t, at a constant percent
solids in the feed, C^, again on logarithmic paper. The slope of this
plot, n, = 0.656, with a standard deviation of 0.0433 and a correlation
coefficient of 0.916. Equation 10 now becomes:
,r 0.767
L = K (11 _ \ (U)
1 0.656 '
68
-------
FIGURE 15
OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE
CAKE YIELD VS. % SOLIDS IN FEED
BASIS:
I0r VACUUM: !5-20"Hg
SUBMERGENCE:25-30%
CYCLE TIME:'V6.3 MIN /REV
FeCI3 ADDED AS % (BY WT) DRY TOTAL SOLIDS: 6-19
FILTRATION AREA: 9.0 FT2
A THICKENED WASTE ACTIVATED SLUDGE
O WASTE ACTIVATED SLUDGE
C\J
o:
o
en
o
UJ
o
'SLOPE = 0.767
CORRELATION COEFFICIENT =0.970
O..U
PERCENT SOLIDS IN FEED
69
-------
FIGURE 16
OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE
CAKE YIELD VS. CYCLE TIME
BASIS:
10
9
e
7
VACUUM: !5-20"Hg
SUBMERGENCE 25-30%
% SOLIDS IN FEED: 2.0-2.4
FeCI3 ADDED AS % (BY WT) DRY TOTAL SOLIDS:8-13
FILTRATION AREA: 9.0 FT2
oc.
*
o 4
o:
o
o
UJ
>-
UJ
-SLOPE =0.656
CORRELATION COEFFICIENT -0.916
_L
2 345
CYCLE TIME (MINUTES/REVOLUTION)
8 9 K)
70
-------
Then constant, K, is obtained from Figure 17 where L is plotted against
on arithmetic paper.
The slope of this plot, K, = 3.85, with a correlation coefficient of
0.90.
Therefore, the cake yield, L, is correlated to be:
(12)
Figure 18 shows the plot of the observed cake yield vs. the calculated
cake yield obtained by using Equation 12 for all runs on oxygenated
waste activated sludge, unthickened as well as thickened, conditioned
with FeCl3, or FeCl3 plus lime. The slope of this plot is 1.0, indica-
ting a good correlation between observed and calculated values. The
standard deviation for the plotted data is 0.22.
The characteristics of the vacuum filtrate are also important since the
filtrate is normally recycled to the aeration bays of a treatment plant.
Tables 1 and 2 in Appendix F provide data on the characteristics of
oxygenated waste sludge filtrate. Typical values derived from these
tables are:
PH = 3.4
TSS (mg/1) = 50.0
VSS (mg/1) =40.0
TDS (mg/1) = 3100.0
VDS (mg/1) = 1000.0
BOD (mg/1) =30.0
The fineness in the weave of the filtration medium used has a major
influence on filtrate quality. The filter medium used in these experi-
ments resulted in excellent solids capture. The recycling of the above
filtrate would impose a minimum organic and solids recycle load on a
secondary treatment system.
COMPARISON OF OXYGENATED AND AIR AERATED WASTE ACTIVATED SLUDGES
The filtration of waste activated sludge is practiced at two major
locations in this country, Milwaukee and Houston. At both of these
locations, waste activated sludge is thickened prior to filtering.
71
-------
5.0r
(M
4.0-
m
FIGURE 17
OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE
OBSERVED CAKE YELD VS.
ELD (OBSERVEC
to c
0 c
L0- 3.85 (f)-0.656 (C.,0.767
^^
.0.767'
0.656
o
o
o
8
o
-SLOPE* 3.85
CORRELATION COEFFICIENT-O.90
0.3
0.4 0.5 0.6
(t)-0.656 . (Cj ) 0.767
0.7
0.8
0.9
1.0
-------
SLOPE =1.0
FIGURE 18
OXYGENATION SYSTEM
WASTE ACTIVATED SLUDGE
CALCULATED VS. OBSERVED CAKE YIELD
O— CONDITIONED WITH FeCI3
T— CONDITIONED WITH FeCI3 a LIME
A— THICKENED SLUDGE CONDITIONED WITH FeCl3
2345
OBSERVED CAKE YIELD (LB /HR /FT2)
-------
Data has been provided by the City of Milwaukee for comparison with the
filtration results obtained at Batavia.
The range of operating conditions of the Milwaukee vacuum filtration
operation during the period when the data were obtained were:
Vacuum: 15-20" Hg
Conditioner: FeCl3 to pH 3.2
Filter (Drum Type): 13' dia. x 16' face
Feed Solids Cone.: 1.4-2.0% thickened waste activated sludge
Cycle Time: 2.0-5.5 rain/revolution
Individual sets of feed solids concentrations and cycle times within this
range of data supplied by Milwaukee were substituted into Equation 12 to
determine the predicted cake yield for oxygenated waste activated sludge,
under the same conditions. These calculated values for the oxygenation
system waste sludge were then compared with the observed Milwaukee cake
yields in Figure 19. From the slope of the plot in Figure 19, it appears
that an average 437., higher yield is possible with unthickened oxygenated
sludge over what is being obtained in Milwaukee with thickened waste
activated sludge. While it must be recognized that the yield enhancement
may be partially attributable to location, it also should be expected
based on the excellent settling and flocculating characteristics of
oxygenated sludge.
ANAEROBICALLY DIGESTED SLUDGE
Table 4 in Appendix F shows the results of pilot-scale runs using the
Batavia plant's anaerobically digested sludge. The objective of these
runs was primarily to compare the performance of the pilot-scale Eimco-
belt filter against the Batavia plant vacuum filter on the same sludge
feed. A filter yield of 5-6 lb/hr/ft2 was obtained on the plant filter
using 57, FeCl-j plus 40% lime (percent of dry solids in feed sludge)
compared with 5.5-9.3 lb/hr/ft2 on the Eimcobelt filter using 5.5-67.,
FeCl3 plus 22-377o lime. It is interesting to note that the optimum chem-
ical condition dosage for this sludge was approximately 57» FeClo plus
approximately 30-40% lime. Outside of this dosage range, no cake was
produced as evidenced by run numbers 2 dnd 3 in Table 4. This is due to
experimental evidence13 that there is an optimum pH for effective condi-
tioning of a particular type of sludge. In this case, the optimum pH was
approximately 12.0. The pH of the conditioned sludge of run numbers 2
and 3 is 8 and 9, respectively, the former due to insufficient lime addi-
tion and the latter due to excess FeCl3 addition (CaO is basic, whereas
FeCl3 is acidic).
74
-------
FIGURE 19
OBSERVED CAKE YIELD FROM VACUUM FILTRATION OF
MILWAUKEE'S CONVENTIONAL THICKENED WASTE ACTIVATED SLUDGE
VS CALCULATED CAKE YIELD OF OXYGENATED SLUDGE
CM
u.
•x.
tr
V.
ffi
O
_l
UJ
>
UJ
u
UJ
O
_l
UJ
UJ
a
UJ
a:
8
SLOPE =0.70
MILWAUKEE FILTER MEDIA
• COTTON CLOTH, WIRE BOUND
• DACRON CLOTH, 16 OZ. PANNEL
A DACRON CLOTH, 8 OZ, WIRE BOUND
1.6 2.0 3.0 4.0
CALCULATED CAKE YIELD BASED ON BATAVIA OXYGENATED SLUDGE (LB/HR/FT2)
-------
AEROBICALLY (OXYGEN AERATED) DIGESTED WASTE ACTIVATED SLUDGE
As part of the aerobic digestion experiments of this contract, the sludge
from the digesters was periodically filtered. The results are tabulated
in Table 5 in Appendix F. It is seen that, in general, the dewatering
characteristics of the aerobically digested sludge were much poorer than
those of the waste activated sludge under the conditions tested.
76
-------
SECTION VIII
AEROBIC DIGESTION
An aerobic digestion study was undertaken at Batavia on oxygenation
system waste activated sludge using high-purity oxygen aeration and
the existing multistage pilot plant assembled under the initial
contract. The digestion experiments were performed entirely on a
batch basis with detention times varying from six to twenty-two days.
Periodically, a portion of the sludge of each run was filtered on the
Eimcobelt pilot-scale vacuum filter.
The results of this study yield the following conclusions concerning
the aerobic digestion of oxygenated waste activated sludge:
1. A non-objectionable stabilized sludge and a 20-307, volatile
suspended solids reduction are obtained after 7-9 days of
digestion. This is comparable with literature values for
air aerated aerobic digestion data under ideal laboratory
conditions .
2. The volatile suspended solids reduction can be represented by
the following equation:
(VSS )
t=t = e-(0.0382t)
(VSS )
1 t=0
where: (VSS ) = total volatile suspended solids at any time,
T t=t t
(VSS ) = total volatile suspended solids at time zero
T t=0
t = detention time (days)
3. Oxygen uptake rate may be expressed as:
02 uptake (mg/l/hr) = 0.0047 (VSS ) - 24.46
1 t=t
4. The vacuum filtration characteristics of aerobically digested sludge
are much poorer than that of waste activated sludge.
Inherent in biological wastewater treatment is the net generation of
excess biological organisms or sludge. The facilities required to
stabilize and dispose of this excess sludge represent a substantial
portion of the cost of wastewater treatment plants. The classical
77
-------
method of sludge disposal, of course, has been anaerobic digestion followed
commonly by vacuum filtration and disposal of the dewatered solids through
landfill, land spreading, or incineration?, in recent years, however, due
to the many problems (including high cost) encountered with the operation
and maintenance of anaerobic digesters, sanitary engineers have begun to
turn to other methods of sludge stabilization. One of these methods is
aerobic digestion.
Aerobic digestion is a process in which micro-organisms obtain energy by
auto-digestion (endogeneous respiration) of the cell protoplasm. The
biologically degradable organic matter in the sludge cells is oxidized
to carbon dioxide, water, and ammonia nitrogen which is further converted
to nitrate and nitrate nitrogen as aerobic digestion proceeds?. The major
advantages of aerobic digestion are that it produces a non-objectionable,
biologically stable sludge suitable for subsequent treatment, and is not
subject to the sensitive operating conditions associated with anaerobic
digestion. Volatile solids reduction similar to anaerobic digestion are
possible with much shorter detention times .
Aerobic digestion with high-purity oxygen would seem to be particularly
compatible with an oxygenation system, since a healthy aerobic biomass
has already been cultivated. The use of oxygen as the aerating gas
would be expected to lend some process advantages. Specifically, the
ease of gas dissolution should make it possible to maintain a higher
dissolved oxygen concentration, which in turn would be expected to
enhance stabilization rates.
The aerobic stabilization experiments were run in a modified oxygena-
tion pilot plant. A new baffle was installed to convert the tankage
from a four-stage, flow-through system to two dual compartment 800-gallon
digesters. The individual digesters were operated batchwise for all runs.
Mixing and aeration were provided by surface aerator-submerged propeller
combinations. Although these units were designed for standard oxygena-
tion pilot plant operation, sufficient excess power was generally avail-
able to provide mixing and oxygen transfer while in sludge digestion
service. The exception occurred each time a digester was refilled with
waste activated sludge. The sludge entered at zero dissolved oxygen
and had an immediate oxygen demand. It then would take 4-6 hours to
produce a dissolved oxygen level greater than 4 mg/1. Due to the
geometry of the tankage, it was necessary to use two surface aerator-
submerged propeller arrangements for each digester. Normal power draw
generally totaled about 1.2 KW per digester.
The aerobic digestion program undertaken here was an initial study.
For this reason, the mechanical components were not designed for
optimum oxygen utilization and continuous operation. Consequently,
high oxygen utilization was not demonstrated in the digestion studies
during this contract work. Oxygen data collected was used solely for
operational control. All the aerobic digestion data (a total of 8
runs) are tabulated in Table 1 of Appendix G.
78
-------
SOLIDS REDUCTION
The results of the aerobic digestion studies indicate that the volatile
solids decrease with detention time and then level off to almost constant
values after 10-15 days, in the endogenous respiration phase, the food
or nutrient level in the environment is insufficient to support the
living biomass, which, therefore, forces the biomass to utilize their
own cell material and surrounding dead cells for food. However, some
of the cell material is not readily biodegradable, i.e., it is consumed
at a negligible ra'te during the detention times used in these aerobic
digestion experiments.
The rate of endogenous decay based on the readily biodegradable material
is represented by the first order reaction:
where : k = rate of decay constant
S = concentration of readily biodegradable cell material at
any time, t, i.e., the readily degradable VSS (mg/1)
t = detention time (days)
Integration of the above equation gives:
ln = -kt
Equation 2 indicates that the plot of the ratio of the readily bio-
degradable VSS at any time, t, to those at zero time vs. detention
time will yield a straight line on semi-logarithmic paper.
Runs No. 6 and 8 were both continued for detention times longer than
ten days. These two runs, shown in Table 2 and 3 in Appendix G, were
analyzed by plotting the ratio of readily biodegradable VSS remaining
each day to those existing intially vs. reactor detention time. Readily
biodegradable VSS were arbitrarily defined as that fraction of the total
VSS consumed during the length of each run. The results of this plot,
shown in Figure 20, can be represented by the following equation:
- (0.12t)
(3>
Also, the ratio of the total volatile suspended solids at any time, t,
to those at time zero is plotted against detention time for all runs as
shown in Figure 21 to yield:
79
-------
i.oc
o
UJ
CO
UJ
01
UJ
o
CO
UJ
s
03
O.I
0.01
FIGURE 20
REDUCTION OF READILY BIODEGRADABLE
VOLATILE SUSPENDED SOLIDS
VS. DETENTION TIME
o RUN NO. 8
o RUN NO. 6
t =e-(O.I2t)
o
8 12 16
DETENTION TIME (DAYS)
20
24
-------
where:
(VSS )
1 t=t
-------
0.7
-T 0.6
r
UJ
M
0.5
O
O
(vssT)t=t
(vssT)t=0
= e-(0.0382t)
v>
£
(/]
w
ui
o
Ul
5
§
0.4
0.3
0.2
FIGURE 21
REDUCTION OF TOTAL VOLATILE SUSPENDED SOLIDS
VS. DETENTION TIME
0.
8 12 16
DETENTION TIME (DAYS)
20
24
-------
e>
o
UJ
tr
LJ
tsl
UJ
0.4
o o
V)
§ 0.3
co
LJ
CO
V)
e 0.2
FIGURE 22
REDUCTION OF TOTAL SUSPENDED SOLIDS
VS. DETENTION TIME
O.I,
10 15 20
DETENTION TIME (DAYS)
25
30
83
-------
I.Or
0.9
0.8
0.7
en
S 0.6
0.5
0.4
0.3
0.2
O.I
O
Q
o
V
O
LEGEND:
RUN NO. I A
RUN NO. 2 •
RUN NO. 3 "3
RUN NO. 4 O
RUN NO. 5 0
RUN NO. 6 ©
RUN NO. 7 •
RUN NO. 8 V
FIGURE 23
RATIO OF VOLATILE SUSPENDED SOLIDS
TO TOTAL SUSPENDED SOLIDS
VS. DETENTION TIME
10 15 20
DETENTION TIME (DAYS)
25
30
-------
of one-third meat packing wastes and two-thirds domestic wastes. A
volatile solids reduction of 35% was reported at 20°C after 8 days
of aeration.
OXYGEN UPTAKE RATES
Oxygen uptake rates were also measured for all the runs as a measure of
the biological activity and oxygen consumption. The instantaneous
oxygen uptake rate under endogenous respiration should be proportional
to the readily biodegradable cell material under aeration at any given
timel7; i.e.,
where
But,
where:
OUR
OUR
b1
S
S
(VSS)
b's (6)
instantaneous oxygen uptake rate (mg/l/hr)
proportionality constant
readily biodegradable VSS at any time, t, (mg/1)
(vssT)
T t=t
(VSS)
SD
(7)
SD
Therefore, OUR
slowly degraded portion of the total VSS
)t=fc - (VSS)SD] (8)
Equation 8 suggests that the instantaneous oxygen uptake rate is linear
with respect to the total volatile suspended solids in the digesters at
any given time. The oxygen uptake rates obtained in the digesters were
plotted against the volatile suspended solids (for all the runs) in
Figure 24. The least squares technique gives:
OUR
0.0047 (VSST) = - 24.46
(9)
with'a correlation coefficient of 0.606. This plot is similar to what
is indicated in the literature and verifies that a rather poor linear
relationship exists for oxygen uptake rates.
NUTRIENT PROFILES
Inspection of the data in Table 1 of Appendix G indicates that aerobic
digestion effected substantial cell destruction during the first 8-10
days of operation as evidenced by the NH3-N and soluble TKN profiles.
Both of the parameters increased from low initial values of less than
50 mg/1 to approach a stable level of 250-300 mg/1 after 8-10 days of
operation. Since the soluble TKN measurement includes ammonia nitrogen,
it is also apparent that the soluble TKN, as measured, consists essen-
tially of ammonia nitrogen and little or no organic nitrogen.
85
-------
O
O
3
•V
UJ
£
cr
Q.
X
O
160
140
120
100
80
60
40
20
FIGURE 24
OXYGEN UPTAKE RATE
VS. VOLATILE SUSPENDED SOLIDS
O
O
OUR =0.0047 (VSST) f=t - 24.46
0
10,000 14,000 18,000 22.0OO 26,000 30.0OO
VOLATILE SUSPENDED SOLIDS (MG/L)
B6
-------
The nitrate and nitrite nitrogen concentrations increase with time dur-
ing each run, although as expected the nitrite nitrogen concentration
never becomes very large since the nitrifying species nitrobactor rapidly
converts nitrite, formed by nitrosomonas, to nitrate nitrogen. In a
plant-scale aerobic digester operation, semi-continuous operation would
normally be maintained as opposed to the batch operation conducted dur-
ing these experiments. Under a semi-continuous operating mode, a popula-
tion of nitrifying bacteria would be maintained within the system so
that cell ammonia would be continuously oxidized as it was released.
In the batch operation conducted during this work, a nitrifying popula-
tion was not available at the start of each run since the Batavia oxygen-
ation system operated at high food-to-biomass loadings exceeding that
necessary to maintain a nitrifying population. Consequently, a long
growth period would be expected in the aerobic digesters for the rela-
tively slow growing nitrifying species. The results of this are seen
as high residual ammonia nitrogen concentrations in each batch run.
Soluble phosphorus concentrations increased during the digestion runs
as shown in Table 1 in Appendix G. This is consistent with cell
destruction and subsequent release of the phosphorus contained in the
cell material.
87
-------
SECTION IX
ACKNOWLEDGEMENTS
This project was conducted under the auspices of the Linde Division's
Process and Product Development Department at Tonawanda, New York
(M. L. Kasbohm - Director). The technical efforts were a part of
the Division's Wastewater Treatment Technology program and involved
direct contributions by the following personnel:
J. R. McWhirter
F. W. Bonnet
E. A. Wilcox (Co-author)
E. K. Robinson
A. W. Bailey
M. A. McDowell (Co-author)
K. W. Young (Co-author)
J. G. Albertsson
N. P. Vahldieck (Co-author)
J. R. Duemmer
S. E. Faruga
E. L. Tytka
D. R. Wyllie
R. R. Tremblay
Manager, Wastewater Treatment
Contract Program Manager
Contract Technology Manager
Consultant
Contract Administrator
Project Engineer
Senior Engineer
Supervisor
Senior Engineer
Technician
Senior Technician
Technologist
Technician
Technician
Appreciation is expressed to the City of Batavia, New York for
permission to conduct the work described here at their Municipal
Pollution Control Plant. The enthusiastic attitude displayed toward
this experimental program by Ira M. Gates, City Manager; Robert
Lawrence, Superintendent of Water and Sewage Works; and his competent
staff contributed greater to the successful completion of the project,
In the course of the project work, the efforts of Richard C. Brenner
(Project Officer), Robert L. Bunch, and James E. Smith, Jr. of the
Advanced Waste Treatment Research Laboratory of the Environmental
Protection Agency were also greatly appreciated.
89
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SECTION X
REFERENCES
1. Albertsson, J. G., McWhirter, J. R., Robinson, E. K., and
Vahldieck, N. P., "Investigation of the Use of High Purity
Oxygen Aeration in the Conventional Activated Sludge Process/1
Federal Water Quality Administration Water Pollution Control
Research Series Publication 17050 DNW, May, 1970.
2. Goackley, P. and Jones, B. P. S., "Vacuum Sludge Filtration.
I. Interpretation of Results by the Concept of Specific
Resistance," Sewage and Industrial Wastes. Vol. 28, No. 8,
pp. 963-975, August, 1956.
3. Poon, C. P. C., "Acceleration of Bio-oxidation with Chemical
Additives and Pure Oxygen," Presented at October, 1969 meeting
of New England WPCA.
4. "Standard Methods for the Examination of Water and Wastewater,"
12th Edition, 1965.
5. Eimco Filter Test Leaf Kit Instruction Booklet, Eimco Corp.,
Salt Lake City, Utah, Publication No. E-1987.
6. "Sludge Dewatering," Manual of Practice No. 20, Water Pollution
Control Federation, Washington, D. C., 1969.
7. Burd, R. S., "A Study of Sludge Handling and Disposal," Federal
Water Pollution Control Administration Water Pollution Control
Research Series Publication No. WP-20-4, May, 1968.
8. Schepman, B. A. and Cornell, C. F., "Fundamental Operating
Variables in Sewage Sludge Filtration," Sewage and Industrial
Wastes. Vol. 28, No. 12, pp. 1443-1460, December, 1956.
9. Walker, W. H., Lewis, W. K., McAdams, W. H., and Gilliand, E. R.,
"Principles of Chemical Engineering," 3rd Edition, McGraw-Hill,
New York and London, 1937.
10. Perry, R. H., et al., Ed. "Chemical Engineers Handbook," 4th
Edition, McGraw-Hill, New York, New York, 1963.
11. Trubnick, E. H., Mueller, P. K., "Sludge Dewatering Practice,"
Sewage and Industrial Wastes, Vol. 30, No. 11, pp. 1364-1378,
November, 1958.
12. Eckenfelder, W. W., Jr., "Water Quality Engineering for Practicing
Engineers," Barnes and Noble, Inc., New York, New York, 1970.
91
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REFERENCES CONTD.
13. Tenney, M. W., Eschelberger, W. F., Jr., Coffey, J. J., and
McAloon, T. J., "Chemical Conditioning of Biological Sludges
for Vacuum Filtration/1 Journal WPCF. Vol. 42, No. 2, Part 2,
pp. R1-R20, February, 1970.
14. Eckenfelder, W. W., Jr., "Studies on the Oxidation Kinetics of
Biological Sludges," Sewage and Industrial Wastes, Vol. 28, No. 8,
pp. 983-989, August, 1956.
15. Jaworski, J., Lawton, G. W., and Rohlich, G. A., "Aerobic Sludge
Digestion," Conference in Biological Waste Treatment, Manhatten
College, April, 1960.
16. Reynolds, T. D., "Aerobic Digestion of Waste Activated Sludge,"
Water and Sewage Works, Vol, 114, No. 2, pp. 37-42, February,
1967-
17. Eckenfelder, W. W., Jr., "Industrial Water Polution Control,"
McGraw-Hill, New York, 1966.
18. Rickard, M. D. and Gaudy, A. F., "Effect of Oxygen Tension on
02 Uptake and Sludge Yield in Completely Mixed Heterogeneous
Populations," Proc. 23rd Purdue Industrial Waste Conference, pp.
883-893, 1968.
92
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SECTION XI
APPENDICES
A.
B.
C.
D.
E.
F.
G.
DESCRIPTION OF BATAVIA PLANT AND EQUIPMENT
FIGURE 1 SCHEMATIC DIAGRAM OF BATAVIA WASTEWATER
TREATMENT PLANT
FIGURE 2 SCHEMATIC DIAGRAM OF A BATAVIA CLARIFIER
FIGURE 3 CROSS-SECTIONAL VIEW OF TYPICAL BATAVIA
OXYGENATION STAGE
FIGURE 4 OVERALL VIEW OF AIR AND OXYGENATION
SYSTEMS AT BATAVIA
FIGURE 5 VIEW OF STAGE NO. 1 OXYGENATION SYSTEM
AERATOR, BLOWER, AND TURBINE DRIVE
METHODS FOR INDIVIDUAL MEASUREMENT
OXYGENATION SYSTEM PERFORMANCE DURING PHASE I
TABLE 1 PHASE I OPERATION - OXYGENATION SYSTEM
PERFORMANCE - WEEKLY AVERAGE OF DAILY VALUES
OXYGENATION SYSTEM PERFORMANCE DURING PHASE II
TABLE 1 PHASE II OPERATION - OXYGENATION SYSTEM
PERFORMANCE - WEEKLY AVERAGE OF DAILY VALUES
OXYGENATION SYSTEM MASS BALANCES
TABLE 1 BATAVIA MASS BALANCES - PHASE I
TABLE 2 AVERAGE OF WEEKLY BATAVIA MASS BALANCES
FOR PHASE I
TABLE 3 BATAVIA MASS BALANCES - PHASE II
TABLE 4 AVERAGE OF WEEKLY BATAVIA MASS BALANCES
FOR PHASE II
VACUUM FILTRATION DATA
TABLE 1 OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE
TABLE 2 THICKENED OXYGENATION SYSTEM WASTE
ACTIVATED SLUDGE
TABLE 3 FILTER LEAF TEST DATA
TABLE 4 ANAEROBICALLY DIGESTED SLUDGE
TABLE 5 AEROBICALLY DIGESTED SLUDGE
TABLE 6 NOMENCLATURE
AEROBIC DIGESTION DATA
TABLE 1 SUMMARY OF DIGESTER DATA
TABLE 2 DIGESTER RUN NO. 6
TABLE 3 DIGESTER RUN NO. 8
Page
95
96
97
99
100
103
105
111
112
117
118
123
124
128
129
131
133
134
146
148
151
153
155
157
158
163
164
93
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APPENDIX A - DESCRIPTION OF BATAVIA PLANT AND EQUIPMENT
The Batavia, New York, Water Pollution Control Plant uses the diffused air
aerated activated sludge treatment process for wastewater treatment. The
plant is designed for a 2.5 MGD average daily flow and a maximum flow of
6.25 MGD. Raw wastewater from the Batavia collection system enters a pump
station located remotely from the treatment plant. Grit and debris are
removed from the wastewater before it is pumped through a 14-inch force
main into two parallel comminutors located at the treatment plant (Point 1
on Figure 1). From the comminutors the wastewater flows by gravity
through a 30-inch main to the aeration tanks. As seen in Figure 1, four
aeration tanks are used and connected hydraulically to form two parallel
treatment systems. The piping of the aeration tanks is such that both
conventional activated sludge or contact stabilization treatment pro-
cesses can be employed by directing raw wastewater through pipe branches
B or C, respectively. The volume of each aeration train (two tanks) is
325,200 gallons, resulting in 6.2 hours of mixed-liquor detention time
based on a 2.5 MGD feed flow. Each tank is 23 feet wide and 64 feet in
length and has a 15'-6" liquid depth.
Figure 2 shows a cross sectional view of the circular clarifiers employed
at Batavia. The clarifiers are designed for center feed and peripheral
takeoff. The side wall depth is 10 feet and the center well depth is
7 feet-10 inches. The total surface area of each clarifier is 1260 ft2
which is equivalent to an overflow rate of 1000 gpd/ft2 in each clarifier
at a design flow of 1.25 MGD to each clarifier. Since the Batavia plant
was designed for operation without primary clarification, a large center
well was installed in each clarifier. Of the total of 1260 ft2 of sur-
face area in each clarifier, 440 ft2 is occupied by the center well. On
this basis, the annular upflow area of each clarifier is only 820 ft2,
so the overflow rate in this section is equivalent to 1500 gpd/ft2 at a
design flow of 1.25 MGD.
During the initial contract, settled activated sludge was pumped from the
clarifiers by two air lift pumps (Point 2 on Figure 1) to a sink (3) from
which the return activated sludge flowed by gravity to the feed end of the
aeration trains. During the second contract, the air lift pumps were
replaced with remotely controlled variable speed centrifugal pumps to
provide a more responsive and wider range sludge return system. Waste
activated sludge also flows by gravity from sink (3) to the sludge
thickener. Return sludge pumping rates for each clarifier can be
independently controlled to meet process conditions. Clarifier piping
is flexible such that both clarifiers can be operated with the same or
separate aeration tanks, a necessary feature for parallel process evalua-
tion. The final treatment step at Batavia is chlorination of the effluent
before it is discharged into Tonawanda Creek.
The multistage high-purity oxygenation process requires that the aeration
95
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FIGURE I-APPENDIX A
SCHEMATIC DIAGRAM OF BATAVIA
WASTEWATER TREATMENT PLANT
BRANCH B
BRANCH C
CHLORINE
CONTACT
TANKS
AERATION TRAIN I
TANK 1
TANK 2
AERATION (OXYGENATION
TRAIN 2
STAGE I STAGE 6
STAGE 2
STAGE 5
TANK 3 | TANK 4
W
STAGE 3 F STAGE 4
r
-o
96
-------
t
10V
40' -I" PI A.
23'-8" DIA.
I
-CENTER
WELL
BAFFLE
\
t
MIXED LIQUOR
TOTAL SURFACE AREA » 1260 SQ.FT.
CENTER FEED WELL SURFACE AREA » 440 SO. FT.
ANNULAR UPFLOW SURFACE AREA • 820 SO. FT.
RETURN SLUDGE
00
o
m
o
c
m
o
1-
d f
O >
TJ
O -o
r- m
£ z
O o
^ x
> X
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-------
tanks be enclosed by a gas-tight cover. Both the gas and liquid phases are
staged and contacted co-currently. These conditions can be readily met
during new aeration tank construction in which both the staging baffles and
the cover can be made an integral structural part of the tanks. At Batavia,
however, the aeration tanks existed and the conversion was a temporary
measure designed for ease of installation and final removal upon comple-
tion of the contract work. For this reason, steel rather than concrete
was chosen as the material of construction for baffles and covers.
Figure 3 is a schematic drawing of one of the oxygenation stages employed
at Batavia. The cover structure for the oxygenation system was designed
for two main functions--to contain the oxygenation gas over the mixed-liquor
at a pressure not exceeding 12 inches of water column and to support the
mixing and gas recirculation equipment. Prefabricated and reinforced
1/8-inch thick steel sheets were used for cover material, each sheet
spanning the tank and overlapping the concrete. A gas-tight seal between
the concrete walls and the steel covers was maintained by an oxygen com-
patible gasketing material. Each cover sheet was compressed against the
gaskets by bolted down clips along all edges of the cover sheet. Access
to the oxygenation stages was provided by 20-inch diameter manways located
in the cover. The manway covers served as safety relief gas vents in the
event of over-pressurization of the system. Relief settings of 8 inches
of water column gauge pressure were maintained by appropriate water-filled
seals in each manway cover.
Sample ports (Item 7, Figure 3) for mixed-liquor sample withdrawal were
provided in each stage. These sample ports extended two feet into the
liquor to prevent gas loss during sampling. Interstage flow of gas was
through 3-inch diameter tubes located above the tank cover. In a perma-
nent installation, interstage transfer of oxygen would be through small
openings in the gas staging baffles beneath the tank covers. An overall
view of the cover and equipment layout is shown in Figure 4. The air aera-
tion tanks are shown in the foreground with the six-stage oxygenation unit
used in the initial contract in the background.
The multistage activated sludge oxygenation process used depends upon
separate mechanical equipment components for liquid mixing and oxygen
compression and dissolution. Both the liquid mixing and the oxygen dis-
solution energy requirements for each stage in a system vary considerably.
Therefore, each stage is equipped with an independent mixer-oxygenator
combination designed to provide the required level of mixing and oxygena-
tion for that particular stage. This arrangement results in efficient
oxygen transfer energy utilization through judicious matching of efficient
mixing and oxygen dissolution equipment to the requirements of each stage
throughout the multistage contacting system. Where large fluctuations in
organic and hydraulic loadings are expected, an automatic dissolved oxygen
level feed back control system is used. This control, which operates by
varying the gas recirculation rate within the stage, may be provided where
desirable.
The oxygenation equipment used at Batavia was designed for efficient opera-
98
-------
VO
VO
FIGURE 3-APPENDIX A
CROSS-SECTIONAL VIEW OF TYPICAL
BATAVIA OXYGENATION STAGE
- VALVE CONTROL
OXYGEN SUPPLY
17-3"
-------
FIGURE 4 - APPENDIX A
OVERALL VIEW OF AIR AND OXYGENATION
SYSTEMS AT BATAVIA
o
o
-------
tion under the wide range of process conditions to be evaluated in the
course of the contract work. A Link-Belt shaft -mounted speed reducer (1)
(size 307D20) provided a double gear reduction of 20:1. This speed
reducer was designed to mount directly onto the shaft to be driven with
the shaft passing through the speed reducer. This feature was necessary
in order to inject oxygen gas into the propeller shaft (4) at Point 10
as shown in Figure 3 .
nn reducer was driven by a 1725-rpm electric motor (2) through
four "B" size v-belts and sheaves. The output, or propeller shaft,
speed when a 1:1 sheave ratio was used was 86.25 rpm. Propeller shaft
speeds of 62.5 and 74.5 rpm were also provided for by selective sheave
ratios. The speed reducer (1), electric motor (2), and propeller
shaft (4) combinations were mounted on a frame (3) designed to integrate
and support the complete oxygenator assembly. The propeller shaft (4)
was a hollow carbon steel tube, 3-1/2 inches in outside diameter with .a
one-half inch thick wall. This shaft supported a three-bladed, 56 inch
diameter marine propeller (6) suspended 126 inches below the tank cover.
Since the overhanging loads at this shaft extension are excessive for the
speed reducer, a steady bearing (5) was designed into the support frame
and located approximately 20 inches below the speed reducer top bearing.
The gas injection sparger device (11) was attached to the propeller shaft
24 inches below the propeller heel center line. The sparger device con-
sisted of a hollow center hub which was threaded to the propeller shaft
and from which 1-1/4 inch diameter pipes extended radially to form the
sparger arms (eight arms per assembly) . Each arm as it extends outward
from the hub was swept back slightly. This swept-back design was suffi-
cient to shed rags or other debris which might collect on the arms as
the sparger rotates. To minimize drag load on the arms, each arm was
tapered by uniformly flattening the arm from 1-1/4 inch outside diameter
at the hub to a 1/2 inch thick, round edged, flat cross section at the
tip. Each arm was drilled through (top and bottom) with 306, 1/8 inch
diameter holes resulting in 2448 holes per sparger.
The compressors (9) used for oxygen gas recirculation in the three stages
were rotary lobe type, manufactured by the Fuller Company - General
American Transportation Corporation, Compton, California. These machines
were designed with dual rotary bronze lobes in a cast iron housing. The
shaft seals at the housing consisted of multiple stationary packing rings
in a gland with a buffer gas introduced in the gland area for pressure
equalization with the process gas, thereby minimizing the process gas
losses .
Feed oxygen gas is fed to the system on demand (a small positive pressure
being maintained by a flow controller to overcome pressure drop through
the unit) with the entire unit operating in effect- as a rc.spirometer .
As the organic load and respiration raLe (oxygen demand) of the biomass
increases, the pressure tends to decrease and feed oxygen flow into the
system increases to re-establish a pressure set point of the controller.
101
-------
Feed oxygen to the multistage system can be controlled on this demand
basis by a simple regulator of differential pressure controller - automatic
valve combination.
An automatic D.O. feedback control and control valve (Point 12 of Figure 3)
was installed in Stage 1 in a bypass arrangement to regulate oxygen flow to
the Stage 1 sparger. The feedback control sensing device was a Union Carbide
Corporation Model 1101 dissolved oxygen probe (13) and analyzer giving a
0 to 50 millivolt linearized output signal. This signal was transmitted
to a millivolt ampere converter (15) having a 1 to 5 ma output (Transmatron
Inc. Model 330T). The converted signal (1-5 ma) was then used by an
indicating controller (8) (Robertshaw Controls Co. Model 321-A1-S2) with
total set point, manual set point, process variable, and valve opening
indications. The 1-5 ma signal was also fed to a 2-inch pipe size auto-
matic diaphragm valve positioner. Proportional conversion of the 1-5 ma
signal to pneumatic activation of the positioner controlled the amount of
bypass gas and consequently the degree of oxygen injection into the mixed-
liquor in Stage 1. The dissolved oxygen control system described was
necessary to meet the specific objectives of the initial contract evalua-
tion program.
The gas seal (14) around the propeller shaft was achieved by two overlapping
cups or cylinders, the lower cup being attached to the cover plate and con-
taining the seal water, the upper cup being attached to the shaft and sealed
to it by an "0" ring. No operational problems were experienced with this
simple hydraulic seal arrangement. A mixed-liquor effluent overflow weir
was installed to prevent rapid changes in tank liquid level, and consequent
changes in gas pressure. This weir also helped prevent bubble entrainment
in the effluent liquor as might occur if the mixed-liquor effluent line
opening were submerged.
A close-up view of a single oxygenator drive, compressor, and automatic
control valve arrangement is shown in Figure 5. It should be pointed out
that these gas-contacting units were designed to permit operation under a
wide variety of process conditions and were highly instrumented. The
device in Figure 5, therefore, does not accurately reflect the type of
equipment which would be employed in a permanent installation.
102
-------
FIGURE 5 - APPENDIX A
VIEW OF STAGE NO. 1 OXYGENATION SYSTEM
AERATOR. BLOWER. AND TURBINE DRIVE
-------
APPENDIX B - METHODS FOR INDIVIDUAL MEASUREMENT
Wastewater Feed Flow
An integrating Fischer-Porter magnetic flow meter was used to monitor
the raw wastewater flow to the oxygenation system. The meter totalizer
was read daily at 10:00 A.M.
Sludge Recycle Flow
An integrating magnetic flow meter in the oxygenation system recycle
sludge line recorded the volume of sludge recycled. Flow integrator
readings were taken daily at 10:00 A.M. as in the case of raw waste-
water flow.
Waste Activated Sludge Flow
An integrating magnetic flow meter in the sludge wasting line recorded
the volume of sludge wasted. Totalized integrator readings were taken
once daily to record the volume of sludge wasted.
Feed Oxygen Flow
A totalizing positive displacement gas flow meter in the feed oxygen line
was read daily at 10:00 A.M., yielding daily total actual volume of feed
oxygen. Average feed oxygen pressures and temperatures were recorded and
used to convert feed oxygen flow rate from actual to NTP (70°F, 14.7 psia)
conditions.
Waste Oxygen Flow
A totalizing positive displacement gas flow meter in the exhaust oxygen
line from the final stage of the system was also read daily as indicated
for feed oxygen. Waste oxygen gas pressures and temperatures were
reporded continuously. The waste gas flow rate at NTP conditions was
calculated from the actual volume of oxygen wasted and the average tem-
perature and pressure.
Gas Recirculation Rate
The rate of oxygen recirculation by each compressor in each stage of
the oxygenation system was measured with an orifice flow meter. The
pressure drop across the orifice, up-stream gas pressure, and recircula-
ting gas temperature were measured and recorded continuously. Oxygen
recirculation rates were calculated from:
105
-------
W = K(Ap • d)1/2
where: W = weight rate of discharge
K = orifice calibration constant
Ap = pressure drop across the orifice
d = density of gas up-stream of the orifice
Power Required for Gas Recirculation
The power required for oxygen recirculation in each stage of the oxygen-
ation system was calculated from the gas recirculation rate, gas inlet
pressure, and compressor discharge pressure by assuming single stage
adiabatic compression at an efficiency of 70 percent.
Oxygen Concentration in Feed, Exhaust, and Stagewise Recirculating Gas
Feed oxygen gas was assumed to be the purity of the temporary liquid
supply from which it was vaporized. Oxygen concentration in the gas
spaces of all oxygenation system stages was continuously monitored
using a multipoint Beckman Model F-3 paramagnetic oxygen analyzer
having an accuracy in this range of + 1/2 percent. Exhaust gas was
assumed to have the same composition as that in the gas space of the
final stage of the system.
Power Required for Liquid Mixing
Power required for liquid mixing in each stage of the oxygenation system
was calculated from wattmeter readings taken daily (for each mixer drive
motor), using motor efficiency values provided by the manufacturer by
assuming a gear-motor efficiency of 85 percent.
Dissolved Oxygen Measurement
The dissolved oxygen concentration in each operating stage of the oxygen-
ation system mixed-liquor, and in the final effluent was monitored con-
tinuously. This was accomplished using a multipoint dissolved oxygen
monitoring system provided by the Instruments Division of Union Carbide
Corporation. In-place probes mounted at a liquid depth of approximately
18 inches were used in each position. The probes were calibrated in
place using a Weston and Stack Model 300E portable dissolved oxygen
meter as .a secondary standard.
Turbidity
This measurement was made on final effluent samples routinely and on
mixed-liquor (settled) supernatant fractions periodically. The in-
strument used was a Hach turbidimeter. The scale calibration was in
106
-------
Jackson Turbidity Units (J.T.U.). Standards were purchased as required.
BOD
The method employed for BOD measurement was that described in "Standard
Methods for the Examination of Water and Wastewater"4. In all instances,
three dilutions per sample were prepared and evaluated. Dissolved oxygen
values before and after incubation were determined using a Weston and
Stack Model 300E Laboratory Dissolved Oxygen Analyzer with the Model A-30
BOD Agitator Assembly, in each instance, the D.O. analyzer was standard-
ized and calibrated against Winkler titration values for replicate
samples. The instrument was highly reliable, as revealed by comparison
tests run during the course of the work, and permitted processing of
significantly more samples than would have been possible if titration
had been required in each instance. All BOD values reported here are
five-day BOD determinations only.
Total and Volatile Suspended Solids and Total and Volatile Dissolved
Solids
These assays were performed on raw wastewater, mixed-liquor, recycle
sludge, and final effluent routinely using essentially the method out-
lined in "Standard Methods"4. The only modification was the use of
glass fiber filters (H. Reeve Angel & Company, Grade 934AH) instead
of gooch crucibles to collect the residue. The tared filter and residue
were dried at 104°C overnight before weighing for total suspended solids.
The same tared filter and residue were then fired at 600°C in a muffle
furnace, cooled in a dessicator, and weighed to permit calculation of
the volatile suspended solids by difference. Blank glass fiber filters
carried through this procedure indicated a weight loss error of only
+0.2 mg., an error well within the sensitivity range of the balance
used in weighing.
Settleability and Settling Rate
This measurement was made using the method given in "Standard Methods" .
The settled volumes at thirty minutes were used for the calculation of
Sludge Volume Index Values. Settled volumes at five minute intervals
from 0 to 30 minutes were recorded. Initial settling rates were calcu-
lated from the five minute interval with the largest change in settled
volume.
Mixed-Liquor Supernatant Analysis
Aliquots of mixed-liquor supernatant after thirty minutes settling in a
one liter cylinder were withdrawn to provide samples for the various
analyses. For mixed-liquor supernatant (non-filtered), this included
COD, BOD, NH3-N, TKN, N02-N, N03-N, and TP.
COD Determination
The method used for COD analysis was basically that described in "Stand-
107
-------
ard Methods"^. In cooperation with the analytical services group of the
Robert A. Taft Water Research Center at Cincinnati during the initial
cpntract, the following modifications were made to facilitate the pro-
cessing of large groups of samples. This method was validated in the
initial contract work by analysis of known COD standards provided by
the Robert A. Taft Center Water Research Center.
A 10 ml sample of raw wastewater, mixed-liquor supernatant, or final
effluent is added to a 50 ml screw cap (Teflon cap) test tube. To this
sample, 20 ml of the standard Ag2S04 solution containing sulfuric acid
is added with mixing. Following this, 10 ml of standardized O.lN K2Cr2Oy
solution containing 0.25 gm of HgS04 is added with mixing. The sample
tube is securely capped and placed in an autoclave where it is digested
at a pressure of 15 psig and a temperature of 120°C for a period of four
hours. After removal of the sample from the autoclave and cooling, the
sample is quantitatively transferred to a titration flask yielding a
final volume with distilled water washing of 200 ml. The COD of the
sample is then determined by back titration of the K2Cr207 using O.lN
Fe(NH4)2(S04)2. The standard ferroin indicator solution is used to
determine the titration end point.
Total Kjeldahl Nitrogen Determination
This analysis was conducted using a Technicon Autoanalyzer. Initial
experimentation, however, indicated the necessity for pre-digestion of
samples containing particulate matter before use of the Technicon system.
To accomplish this, 20 ml samples of raw wastewater, final effluent, or
mixed-liquor supernatant were placed in 50 ml volume screw cap tubes
(Teflon cap) and 20 ml of a 337» H2S04 solution was added to each with
mixing. The sealed sample tubes were then autoclaved at 15 psig and
120°C for 4 hours. This pre-digestion step completely solubilized all
particulate material assuring the admission of a homogeneous sample to
the Technicon Analyzer.
The pre-digested samples were then analyzed for total Kjeldahl nitrogen
using the Technicon Autoanalyzer apparatus with the method described in
Technicon Autoanalyzer Methodology, Bulletin N-3c. The only modification
of the method as described was the addition of 3.570 HgCl2 solution to
the sample prior to passage through the spiral digestion tube. This
modification was suggested by the Robert A. Taft Water Research Center
in addition to the H2S04, Se02, and perchloric acid mixture normally
used as necessary for complete conversion of organic nitrogen to ammonia
nitrogen. Accuracy of the method described was verified by analyzing
known TKN standard solutions.
Ammonia Nitrogen Analysis
Analysis of wastewater samples for ammonia nitrogen was done using the
Technicon Autoanalyzer. The method used was that described in Technicon
Bulletin N-3c for TKN analysis except that the digestion step was elimi-
nated and the sample was treated and analyzed directly.
108
-------
Total and Ortho-phosphorus Analyses
The method used for total phosphorus analysis was essentially that
described in "Standard Methods"^, the Aminoaptholsulfamic Acid Method.
The uniformity of the wastewater sample was assured by digesting it at
15 psig and 120°C in an autoclave for four hours in the presence of
sulfuric acid and ammonium persulfate. Subsequent segments of the ana-
lytical procedure were carried out using a Technicon Autoanalyzer. This
procedure essentially ensured conversion of all organically bound phos-
phorus to ortho-phosphorus. For ortho-phosphorus analysis, the sample
was next treated with acid ammonium molybdate to form molybdophosphoric
acid which yields an intense blue color complex upon the addition of a
solution containing l-amino-2-napthol-4-sulfamic acid, sodium bisulfite,
and sodium sulfite. The blue color complex is developed by heating at
95°C and is measured spectrophotometrically at 660my,.
Nitrite and Nitrate Nitrogen Analyses
These analyses were carried out through an automated procedure using the
Technicon Autoanalyzer. The automated method is essentially the Diazo-
tization Method described in the Environmental Protection Agency Training
Manual for Laboratory Analysis in treatment plant operation. For nitrite
nitrogen analysis, the sample is treated directly with the sulfanilic
acid and a-napthylamine to form an intense red azo dye. The color in-
tensity was measured spectrophotometrically at 520 m|j,.
Analysis for nitrate nitrogen in the sample employs the foregoing pro-
cedure after the nitrate is reduced to nitrite by treatment with hydra-
zine sulfate in an alkaline medium heated to 53°C. Nitrate nitrogen is
measured as the difference between values obtained for nitrite nitrogen
alone and reduced nitrate nitrogen plus nitrite nitrogen.
109
-------
APPENDIX C - OXYGENATION SYSTEM PERFORMANCE DURING PHASE I
111
-------
lABLt I - APPE.VDIX C
WEEKLY AVF.RAfiE OF DATL.Y VALUES
PARAMETER
Raw Wastewater Feed Rate (MGD)
Sludge Recycle Rate (MGD)
'/. Sludge Recycle
Raw Wastewater Temperature (,°F)
Average Mixed-Liquor D.O. (mg/1)
Clarifier Etfluent D.O. (rug/1)
Aeration Detention time (hr)
(Wastewater + Recycle)
Nominal Aeration Detention Time (hr)
(Wastewater Only)
Raw Wastewdter pH
Mixed-Liquor pH - Inlet Stage 1
Middle Stage 2
Outlet Stage 3
8/30/70
3.18
1 .09
34
69
7.4
1.6
0.91
1.23
7.2
6.6
6.5
6.5
9/6/70
2.48
1.86
35
69
8.0
1.2
1.17
1.57
7.0
6.5
6.5
6.3
9/13/70
2.55
0.93
36
68
8.9
1.2
1.12
1.53
6.9
6.5
6.5
6.4
9/20/70
2.91
1.16
40
68
5.6
0.5
0.96
1.34
7.3
6.6
6.6
6.5
9/27/70
2.76
1.48
54
68
6.6
0.7
0.92
1.41
7.5
7.0
6.9
6.8
10/4/70
2.76
1.64
59
68
9.9
0.7
0.89
1.41
7.5
7.2
7.0
6.9
10/11/70
2.82
1.19
42
66
9.0
0.8
0.97
1.39
7.6
7.1
7.0
6.9
10/18/70
2.89
1.17
41
67
10.2
<0.5
0.96
1.35
7.5
6.6
6.5
6.4
AVG.
2.79
1.19
43
68
8.2
0.8
0.98
1.40
7.3
6.8
6.7
6.6
Clarifier Effluent pH
6.3
6.2
6.2
6.5
6.7
6.7
6.6
6.5
6.5
-------
TABLE I - APPENDIX C CONTD.
Food/Biomass Loading (Ib BOD applied/lb
MLVSS)
Volumetric Organic Loading (Ib BOD applied/
day/1000 ft3 tnixed-liquor)
Clarifier Surface Overflow Rate
(gpd/ft*)*
Clarifier Mass Loading Rate (Ib TSS/
day/ft2)
Sludge Settling Rate (ft/hr)**
Sludge Volume Index
Dry Solids Wasted (Ib/day)
Lb VSS Wasted/Lb BOD Removed
Total Feed 02 Flow Rate (cfm at NTP)
Average 7= 02 Composition in Gas Space
of each Oxygenation System Stage
Stage 1
Stage 2
Stage 3
Exhaust Gas Rate from 02 System,
3rd Stage (cfm at NTP)
Overall 7, of Feed Oxygen Utilized
Lb 02 Utilized/ib BOD Consumed
Lb 02 Dissolved/(HP-hr) for Gas Recirculation
and Liquid Mixing
8/30
1.16
272
1260
75.7
4.2
52
4800
0.54
-
84
77
66
2.9
-
-
9/6
0.99
221
984
60.7
4.5
54
2900
0.47
50.6
85
79
69
3.6
95.1
1.26
9/13
1.23
271
1010
50.5
5.4
50
4920
0.57
41.8
83
75
58
5.6
92.2
0.82
9/20
0.74
181
1160
70.2
6.3
51
3200
0.64
43.3
81
68
49
6.7
92.4
1.33
9/27
0.84
190
1110
76.7
4.6
52
3180
0.60
40.5
80
70
51
6.1
92.3
1.15
10/4
0.91
214
1110
75.8
4.2
59
2360
0.41
39.4
81
69
51
5.4
93.1
1.02
10/11
1.10
239
1120
64.8
5.0
55
3020
0.48
41.2
82
70
52
4.0
95.0
0.97
10/18
1.12
260
1150
75.9
5.1
58
3580
0.51
42.5
82
71
54
4.7
94.0
0.91
AVG.
1.01
231
1110
68.8
4.9
54
3490
0.52
42.8
82
72
56
4.9
93.4
1.01
8.78 6.94 7.17 6.68 6.65 6.96
7.13
7.19
* Based on total clarifier surface area of 1260 ft2.
** Based on five minute interval settled volumes in a
1000 ml graduate without stirri.
-------
TABLE I - APPENDIX C CONTD.
8/30 9/6
9/13
9/20
9/27
10/4
10/11
10/18
AVG.
Raw Wastevacer Feed (mg/1)
TSS
IDS
VSS
VDS
VSS/TSS
Mixed-Liquor (tng/1)
(Average all stages)
TSS
IDS
VSS
VDS
VSS/TSS
Clarifier Effluent ( mg/1 )
TSS
TDS
VSS
VDS
VSS/TSS
Recycle Sludge (ng/1 )
TSS
TDS
VSS
VDS
VSS/TSS
134
640
107
154
0.80
5610
575
3770
125
0.67
53
559
41
121
0.77
23, 700
738
14,900
199
0.63
235
664
193
161
0.82
4.820
637
3600
154
0.75
49
623
41
116
0.84
26,460
824
18,200
242
0.69
230
598
177
163
0.77
4720
554
3520
123
0.75
25
563
18
104
0.72
21,000
617
14,500
131
0.69
126
472
100
120
0.79
5300
575
3900
108
0.74
15
516
10
64
0.67
24,000
881
16,500
302
0.69
137
603
97
131
0.71
5030
664
3610
121
0.72
26
614
21
113
0.81
25,800
677
18,400
169
0.71
91
681
71
193
0.78
5320
669
3780
113
0.71
43
644
40
129
0.93
15,000
684
10, 700
88
0.71
142
661
109
148
0.77
4810
727
3490
157
0.73
26
627
19
85
0.73
19,800
736
13,600
152
0.69
130
579
92
143
0.71
5080
978
3710
404
0.73
20
516
16
95
0.80
16,900
780
12,500
118
0.74
153
612
118
152
0.77
5080
672
3680
163
0.72
32
583
26
103
0.79
21,600
742
14,900
175
0.69
-------
TABLE I - APPENDIX C CONTD.
8/30 9/6 9/13 9/20 9/27 10/4
10/11
10/18 AVG.
Ra* Wastevater Feed (mg/1)
BOD
COD
Ortho-phosphorus as P
Total Phosphorus as P
NH3-N as N
TKN as N
N02-N as N
N03-N as N
Clarifier Effluent (mg/1 )
BOD
Soluble BOD
COD
Soluble COD
Ortho-phosphorus as P
Soluble Ortho-phosphorus as P
Total Phosphorus as P
Total Soluble Phosphorus as P
NH3-N as N
Soluble NHj-N as N
TKN as N
Soluble TKN as N
N02-N as N
Soluble N02-N as N
N03-N as N
Soluble X03-N as N
233
414
9.4
•10.8
20.4
33.9
0
0
27
9
67
45
10.3
8.7
9.9
8.3
18.1
13.7
23.8
17.6
0.3
0.4
2.6
2.1
232
524
11.5
14.8
24.7
38.4
0
0.1
55
13
105
59
10.2
10.1
10.9
9.9
16.8
15.6
24.2
20.4
1.1
0.9
3.6
3.6
277
477
9.2
13.6
23.4
35.0
0.7
0.6
51
14
77
59
9.9
9.9
10.1
9.7
19.3
19.4
25.3
23.3
1.1
1.0
3.2
3.2
162
363
7.7
11.5
16.4
29.0
0.4
0.9
30
14
73
60
10.1
10.0
9.9
9.3
15.2
15.3
21.2
18.0
1.0
1.0
1.6
1.6
179
285
9.4
13.7
23.4
33.9
0.8
1.3
35
11
87
86
12.4
12.3
13.4
13.1
22.1
21.5
23.6
24.1
0.7
0.7
4.6
4.6
202
336
8.6
11.8
18.8
30.4
0.2
0.6
42
16
70
62
10.7
10.6
11.1
11.0
16.5
15.9
19.8
18.7
0.3
0.4
4.8
5.4
221
378
7.8
12.2
26.5
42.6
0.6
1.4
41
16
-
-
10.5
10.5
10.0
10.2
24.9
24.0
32.1
28.7
1.3
1.4
2.3
2.3
235
395
7.4
12.0
19.5
36.0
0.4
1.4
42
16
95
55
10.7
11.1
11.3
11.6
18.6
20.8
23.8
20.4
0.8
0.7
3.4
3.2
218
397
8.9
12.6
21.6
34.9
0.4
0.8
40
14
82
61
10.6
10.4
10.8
10.4
18.9
18.3
24.2
21.4
0.8
0.8
3.3
3.2
-------
APPENDIX D - OXYGENATION SYSTEM PERFORMANCE DURING PHASE II
117
-------
TABLE 1 - APPENDIX D
PHASE II OPERATION - OXYGENATION SYSTEM PERFORMANCE
PARAMETER
Raw Wastewater Feed Rate (MGD)
Sludge Recycle Rate (MGD)
7. Sludge Recycle
Raw Wastewater Temperature (°F)
Average Mixed-Liquor D.O. (mg/1)
Clarifier Effluent D.O. (mg/1)
Aeration Detention time (hr)
(Wastewater + Recycle)
Nominal Aeration Detention Time (hr)
(Wasteuater Only)
Raw Wastewater pH
Mixed-Liquor pH - Inlet Stage 1
Middle Stage 2
Outlet Stage 3
WEEKLY AVERAGE
11/1/70
1.95
0.62
32
65
) 10.6
0.8
1.52
(hr)
2.00
7.1
1 6.5
> 2 6.4
: 3 6.3
OF DAILY VALUES
11/8/70
2.19
0.70
32
63
13.1
<0.5
1.35
1.78
7.0
6.4
6.3
6.2
11/15/70
2.19
0.76
35
61
12.5
<0.5
1.32
1.78
7.0
6.5
6.5
6.5
11/22/70
2.22
1.00
45
60
12.4
<0.5
1.21
1.76
7.2
6.8
6.6
6.6
AVG
2.14
0.77
36
62
12.1
<0.5
1.34
1.82
7.1
6.5
6.4
6.4
Clarifier Effluent pH
6.5
6.3
6.5
6.6
6.5
-------
TABLE 1 - APEENDLK D CONTD.
11/1/70 11/8/70 11/15/70 11/22/70 AVG.
Food/Biomass Loading (Ib BOD applied/day/ 0.45 0.66 0.69 0.71 0.63
Ib MLVSS)
Volumetric Organic Loading (Ib BOD applied/
day/1000 ft3mixed-liquor) 73 106 103 121 101
Clarifier Surface Overflow Rate
(gpd/ft2)* 1550 1740 1740 1760 1700
Clarifier Mass Loading Rate
(Ib TSS/day/ft2) 63.8 60.7 75.4 92.0 73.0
Sludge Settling Rate (ft/hr)** 8.9 8.0 8.7 6.8 8.1
Sludge Volume Index 48 56 47 48 50
Dry Solids Wasted (Ib/day) 1150 1930 1260 1470 1450
Lb VSS Wasted/Lb BOD Removed 0.52 0.66 0.37 0.46 0.50
Total Feed 02 Flow Rate (cfm at NTP) 28.6 30.2 28.8 24.6 28.0
Average % 02 Composition in Gas Space
of each Oxygenation System Stage
Stage 1 87 81 82 81 83
Stage 2 73 71 73 72 72
Stage 3 63 59 61 61 61
Exhaust Gas Rate from 02 System,
3rd Stage (cfm at NTP) 4.2 4.3 4.0 3.6 4.0
Overall % af Feed Oxygen Utilized 90.8 91.5 91.6 91.1 91.3
Lb 02 Utilized/Lb BOD Consumed 2.18 1.52 1.50 1.10 1.50
Lb 02 Dissolved/(HP-hr) for Gas Recirculation
and liquid Mixing 5.2 6.4 5.8 - 5.8
* Based on noted clarifier surface area of 1260 ft2.
*-* Based on five minute interval settled volumes in a 1000 ml graduate cylinder without stirring.
-------
TABLE 1 - APPENDIX D CONTD.
11/8/70
11/15/70
11/22/70
AVG.
Rav Wastewater Feed (mg/l)
TSS
TDS
VSS
VDS
VSS/TSS
Mixed-Liquor (mg/1)
TSS
TDS
VSS
VDS
VSS/TSS
Clarifier Effluent (mg/1)
TSS
TDS
VSS
VDS
VSS/TSS
Recycle Sludge (mg/1)
TSS
TDS
VSS
VDS
VSS/ TSS
104
715
87
115
0.84
3640
718
2000
120
0.72
23
687
18
116
0.78
26,700
823
15,500
191
0.68
92
603
76
143
0.83
3440
648
2560
117
0.75
48
537
39
83
0.81
21,700
659
13,700
176
0.63
66
660
52
72
0.79
3560
727
2400
164
0.67
68
639
54
85
0.79
24,300
827
13,700
272
0.57
68
666
37
81
0.54
3980
804
2740
144
0.69
184
647
66
66
0.36
15,600
826
10,500
211
0.67
83
661
63
103
0.76
3650
724
2580
136
0.71
81
627
44
87
0.69
21,100
784
13,300
212
0.63
-------
TABUS 1 - APPENDIX D CONTD.
11/15/70 11/22/70 AVG.
Raw Wastewater Feed (mg/1)
BOD
COD
Ortho-phosphorus as P
Total Phosphorus as P
NH3-N as N
TKN as N
N02-N as N
N03-N as N
Clarifier Effluent (mg/1)
BOD
Soluble BOD
COD
Soluble COD
Ortho-phosphorus as P
Soluble Ortho Phosphorus as P
Total Phosphorus as P
Total Soluble Phosphorus as P
NH3-N as N
Soluble NH3-N as N
TKN as N
Soluble TKN as N
N02-N as N
Soluble N02-N as N
N03-N as N
Soluble N03-N as N
97
194
3.1
7.2
9.3
19.1
0.6
2.5
30
10
86
51
5.3
5.2
6.6
5.9
12.6
12.1
17.5
16.5
0.4
0.4
3.1
3.2
126
195
4.1
7.3
10.0
18.8
0.4
2.3
51
8
78
16
6.1
6.2
7.4
6.4
12.2
11.8
18.2
13.2
0.4
0.3
3.7
3.8
122
149
3.3
6.7
7.9
15.8
0.5
4.2
34
7
112
41
4.3
4.5
7.3 •
4.8
11.6
10.5
19.1
14.6
0.6
0.6
5.5
6.1
142
170
4.4
6.5
12.8
21.0
0.4
1.8
48
9
111
-
5.4
4.3
6.5
6.9
11.1
6.5
17.7
16.9
0.6
0.3
4.7
2.7
122
177
3.7
6.9
10.0
18.7
0.5
2.7
41
8.5
97
36
5.3
5.0
6.9
6.0
11.9
10.2
18.1
15.3
0.5
0.4
4.2
3.9
-------
APPENDIX E - OXYGENATION SYSTEM MASS BALANCES
123
-------
TABLE 1 - APPENDIX E
BAIAVIA MASS BALANCES - PHASE I
Pounds Per Day
WEEK OF
9-6-70
TSS
vss
TDS
VDS
TS
VS
IS
AC
.,02
•M2
Mass
Feed Raw
Oxygen Wastevater
4564
3748
12896
3127
17460
6875
10585
6048
6048 17460
Digester
Supernatant
472
327
39
19
511
346
165
511
Pounds Per Day
WEEK OF
9-13-70
TSS
VSS
TDS
VDS
TS
VS
IS
AC
AO,
AH,
Mass
Feed Raw
Oxygen Wastewater
4577
3522
11901
3244
16478
6766
9712
4968
4968 16478
Digester
Supernatant
444
269
49
23
493
292
201
493
In
Vacuum
Filtrate
75
28
51
10
126
38
88
126
In
Vacuum
Filtrate
220
83
150
28
370
111
259
370
C12 Tank
Transfer
121
22
37
7
158
29
129
158
C12 Tank
Transfer
119
21
34
6
153
27
126
153
Total
Ib/day In
5232
4125
13023
3163
18255
7288
10967
24303
Total
Ib/day In
5360
3895
12134
3301
17494
7196
10298
22462
Pounds Per
Waste Clarifier
Oxygen Effluent
1009
845
12833
2389
13842
3234
10608
20 1171
349 3922
158
369 19093
Pounds Per
Waste Clarifier
Oxygen Effluent
525
378
11834
2186
12359
2564
9795
30 1747
462 5839
235
492 20180
Day Out
Waste
Sludge
3028
2280
92
26
3120
2306
814
3120
Day Out
Waste
Sludge
4971
3141
115
25
5086
3166
1920
5086
Total
Ib/day Out
4037
3125
12925
2415
16962
5540
11422
22582
Total
Ib/day Out
5496
3519
11949
2211
17445
5730
11715
25758
Out
In
0.77
0.76
0.99
0.76
0.93
0.76
1.04
0.93
Out
In
1.03
0.90
0.98
0.67
1.00
0.80
1.14
1.15
Mass In = TS. + AO,
in z
Mass Out = TS + AC + AO, + H,
out out out 'out
-------
TABLE 1 - APPENDIX E CONTD.
Pounds Per Day In
Pounds Per Day Out
WEEK OF
9-20-70
TSS
VSS
TDS
VDS
TS
VS
IS
AC
A02
Olt
Mass
WEEK OF
9-27-70
TSS
VSS
TDS
VDS
TS
VS
IS
AC
•&2
AH2
Mass
Feed Raw Digester Vacuum
Oxygen Wastewater Supernatant Filtrate
3217
2551
12431
3171
15648
5722
9926
5160
5160 15648
486
300
62
33
548
333
215
548
Pounds Per Day
Feed Raw Digester
Oxygen Wastewater Supernatant
2985 600 *
2113
13136
2854
16121
4967
11154
4824
4824 16121
Mass In = TS
237
51
15
651
252
399
651
in 2in
70
26
48
9
118
35
83
118
In
Vacuum
Filtrate
73
27
50
9
123
36
87
123
C12 Tank
Transfer
119
21
31
4
150
25
125
150
C12 Tank
Transfer
120
21
37
7
157
28
129
157
Mass Out
Total
Ib/day In
3892
2898
12572
3217
16464
6115
10349
21624
Total
Ib/day In
3778
2398
13274
2885
17052
5283
11769
21876
= TS
out
Waste Clarifier
.Oxygen Effluent
362
241
12451
1544
12813
1785
11028
31 1211
475 4048
164
506 18236
Pound s Per
Waste Clarifier
Oxygen Effluent
586
481
14069
2589
14665
3070
11595
13 607
405 2038
82
418 17392
*• AC ,. + AO, .. + A!
out out
Waste
Sludge
3192
2287
98
22
3290
2309
981
3290
Day Out
Waste
Sludge
2997
2179
132
25
3129
2204
925
3129
Total
Ib/day Out
3554
2528
12549
1566
16103
4094
12009
22032
Total
Ib/day Out
3593
2660
14201
2614
17794
5274
12520
20939
Out
In
0.91
0.87
1.00
0.49
0.98
0.67
1.16
1.02
Out
In
0.95
1.11
1.07
0.91
1.04
1.00
1.06
0.96
-------
TABLE I - APPENDIX E CONTD.
Pounds Per Day In
Pounds Per Day Out
WEEK OF
10-4-70
TSS
VSS
TDS
VDS
TS
VS
IS
^C
^0,
£»t
Mass
Feed Raw
Oxygen Was tew;
1983
1547
14836
4205
16819
5752
11067
4704
4704 16819
Digester
iter Supernatant
757
478
44
21
801
499
302
801
Vacuuir
Filtrate
45
25
31
8
76
33
43
76
C12 Tank
Transfer
121
22
38
8
159
30
129
159
Total
Ib/day In
2906
2072
14949
4242
17855
6314
11541
22559
Pounds Per Day In
WEEK OF
10-11-70
TSS
VSS
TDS
VDS
TS
VS
IS
-------
TABLE 1 - APPENDIX E CONTD.
Pounds Per Day In
Pounds Per Day Out
WEEK OF
10-18-70
TSS
VSS
IDS
VDS
TS
VS
IS
J\C
A02
AH2
Mass
Feed Raw
Oxygen Wastevater
2964
2098
14729
3260
17693
5358
12335
5064
5064 17693
Digester
Supernatant
607
401
50
25
657
426
231
657
Vacuum
Filtrate
384
154
138
27
522
181
341
522
C12 Tank
Transfer
119
21
36
6
155
27
128
155
Total
Ib/day In
4074
2674
14917
3318
19027
5992
13035
24091
Waste Clarifier
Oxygen Effluent
478
383
14488
2271
14966
2654
12312
22 792
361 2641
108
383 18507
Waste
Sludge
3577
2699
167
45
3744
2744
1000
3744
Total
Ib/day Out
4055
3082
14655
2316
18710
5398
13312
22634
Out
In
1.00
1.15
0.98
0.70
0.98
0.90
1.02
0.94
Mass In = TS. + AO,
in *
Mass Out
TS
out
AC
out
W + ^'out
-------
TABLE 2 - APPENDIX E
AVERAGE OF WEEKLY BAIAVIA MASS BALANCES FOR PHASE I
TSS
VSS
IDS
VDS
TS
VS
IS
AC
•*>*
AH,
Mass
Pounds Per Day In
Feed
Oxygen
5098
5098
Raw
Wastewater
3350
2572
13,518
3307
16,867
5879
10,988
16,867
Digester
Supernatant
537
330
50
23
587
353
234
587
Vacuum
Filtrate
157
62
78
16
235
77
158
235
C12 Tank
Transfer
120
21
36
6
156
27
128
156
Total
Ib/day In
4164
2985
13,682
3352
17,845
6336
11,508
5098
22,943
Pounds Per Day Out
Waste Clarifier
Oxygen Effluent
652
527
13,577
2274
14,229
2800
11,429
19 989
381 3307
133
400 18,658
Waste
Sludge
3306
2381
134
29
3441
2410
1030
3441
Total
Ib/day Out
3958
2908
13,711
2303
17,670
5210
12,459
1008
3688
133
22,499
Out
In
0.95
0.97
1.00
0.69
0.99
0.82
1.08
0.72
0.98
-------
TABLE 3 - APPENDIX E
BATAVIA MASS BALANCES - PHASE II
WEEK OF
11-1-70
TSS
VSS
TDS
VDS
TS
VS
IS
AC
A02
AH2
Mass
WEEK OF
11-8-70
TSS
VSS
TDS
VDS
TS
VS
IS
J£
-°2
j#2
Mass
Feed Raw
Oxygen Wastewater
1704
1426
11723
1885
13428
3311
10117
3408
3408 13428
Feed Raw
Oxygen Wastewater
1645
1359
10781
2557
12435
3916
8519
3600
3600 12435
Pounds Per Day In
Digester Vacuum
Supernatant Filtrate
806
517
54
21
860
538
322
860 0
Pounds Per Day In
Digester Vacuum
Supernatant Filtrate
1165
667
85
49
1250
716
534
1250 0
Cl, Tank
Transfer
60
10
18
3
78
13
65
78
C12 Tank
Transfer
60
10
18
3
78
13
65
78
Total
Ib/day In
2571
1953
11795
1909
14366
3862
10504
17774
Total
Ib/day In
2870
2036
10884
2609
13763
4645
9118
17363
Pounds Per
Waste Clarifier
Oxygen Effluent
381
298
11379
1921
11760
2219
9541
28 715
420 2402
98
448 14975
Pounds Per
Waste Clarifier
Oxygen Effluent
870
707
9735
1505
10605
2212
8341
39 1124
445 3759
154
484 15642
Day Out
Waste
Sludge
1153
745
42
9
1195
745
450
1195
Day Out
Waste
Sludge
1926
1433
64
12
1990
1445
545
1990
Total
Ib/day Out
1534
1043
11421
1930
12955
2964
9991
16618
Total
Ib/day Out
2796
2140
9799
1517
12595
3657
8936
18116
Out
In
0.60
0.54
0.97
1.01
0.90
0.76
0.95
0.93
Out
In
0.97
1.05
0.90
0.58
0.91
0.79
0.98
1.04
Mass In = TS
in
Mass Out
TS
-------
TABLE 3 - APPENDIX E CONTD.
Pounds Per Day In
Founds Per Day Out
ICEK OF
11-15-70
TSS
VSS
TDS
YDS
TS
VS
IS
£C
AO,
£H,
Mass
Feed Raw
Oxygen Wastewater
1188
936
11883
1296
13071
2232
10839
3432
3432 13071
Digester Vacuum
Supernatant Filtrate
965
681
61
44
1026
725
301
1026 0
Clj Tank
Transfer
60
10
18
3
78
13
65
78
Total
Ib/day In
2213
1627
11962
1343
14175
2970
11205
17607
Waste Clarifier
Oxygen Effluent
1237
982
11621
1546
12858
2528
10330
21 677
366 2260
92
387 15887
Waste
Sludge
1260
781
38
8
1298
789
509
1298
Total
Lb/day Out
2497
1763
11659
1554
14156
3317
10839
17572
Out
In
1.13
1.08
0.97
1.16
1.00
1.12
0.97
1.00
Mass In
TSin+A°'in
Mass Out
TS
-------
TABLE 4 - APPENDIX E
AVERAGE OF WEEKLY BATAVIA MASS BALANCES FOR PHASE II
TSS
VSS
TDS
VDS
TS
VS
IS
AC
AOj
AH2
Mass
Feed Raw
Oxygen Wastewater
1513
1240
11,462
1913
12,978
3153
9825
3480
3480 12,978
Founds Per Day In
Digester Vacuum
Supernatant Filtrate
979
622
67
38
1045
660
386
1045 0
Pounds Per Day Out
C12 Tank
Transfer
60
10
18
3
78
13
65
78
Total
Ib/day In
2552
1872
11,547
1954
14,101
3826
10,276
3480
17,581
Waste Clarifier
Oxygen Effluent
829
662
10,912
1657
11,741
2320
9404
29 839
410 2807
115
439 15,502
Waste
Sludge
1446
986
48
10
1494
993
501
1494
Total
Ib/day Out
2275
1648
10,960
1667
13,235
3313
9905
868
3217
115
17,435
Out
In
0.89
0.88
0.95
0.85
0.94
0.87
0.96
0.92
0.99
Mass In = TS.n
Mass Out = IS
out
AO, + AH,
2out 2out
-------
APPENDIX F - VACUUM FILTRATION DATA
133
-------
TABLE 1- APPENDIX F
OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE
RUN NUMBER
DATE
Filter Flow Rate (eptn)
Drum Speed (fc/min.)
Chemical Addition
Line (20* Solution) (ml/mio)
FeCl,(l07. Solution) (ml/mini
pH of Feed Afcer Chemical Add.
Feed Sludge
TSS (mR/n
VSS (n.R/1)
TDS (m*/l)
VDS Og/l)
TS • TSS + TDS (mg/1)
Running Time, (mm)
7. Solids in Cake
Filtrate
pU
TSS (mg/1)
"VSS (mg/1)
TDS (mg/1)
VDS (mg/1)
TS - TSS + IDS (mg/1)
BOD (mg/1)
HHj-N (mg/1)
H02-N (mg/1)
H03 -N (ma/U
1
9/1/70
1.6
1.49
136
136
12.0
26,632
.4,332
840
192
!7,^72
30.0
24.9
11.8
228
104
2,414
716
2,642
2
9/1/70
1.6
1.49
100
136
10.4
26,632
14,332
840
192
27,472
33.6
24.1
9.9
314
122
3,708
1,408
4,022
12.4
3
9/1/70
1.6
1.49
136
4,2
26,632
14,332
840
192
27,472
30.0
20.5
5.4
174
74
2,902
900
3.076
26.9
4
9/1/70
1.6
1.49
..
5.9
26,632
14,332
840
192
27,472
30.0
17.1
6.9
134
104
1,124
356
1,258
22.6
5
9/2/70
1.6
1.49
138
138
H.8
23,456
14,584
656
244
24,112
30.0
23.7
11.6
146
98
2,838
684
2,984
10.2
6
9/2/70
1.6
1.49
138
12.1
23,456
14,584
656
244
24,112
15.0
18.8
11.8
190
164
2,022
616
2.212
8.0
7
9/9/70
1.6
1.49
138
260
9.5
23,504
16,202
342
132
23,846
30.0
17.25
8.7
84
55
4,322
1,298
4,406
16
8
9/9/70
1.6
1.49
136
130
11.6
23,504
16,202
342
132
23,846
30.0
20.5
11.3
42
39
2.894
780
2.936
18
9
9/9/70
1.6
1.49
65
135
7.0
23.504
16.202
342
132
23,846
17.0
15.4
9.6
225
82
1.858
646
2.083
14.0
-------
TABLE I - APPENDIX F COJTD.
RUN NUMBER
DATE
TKH ("*/!)
TP fme/n
ORTHO-P fnu>m
Cycle Time (min/rev)
I (bT wt.) Solids to Filter
Chemical Dosage (as I by wt.
of dry solids in sludge to
filter)
Lime-
Fed,
Polymer
Filtrated Collected (gal /hr)
Filter Cake Moisture
(Ib H,0/lb dry solids)
Cake Yield
(Ib dry solids/hr/ft2) .
Remarks
1
A/70
6.3
2.75
18.5
8.7
89.4
3.02
2.00
Good
Cake
2
9/1/70
27.8
2.5
...
6.3
2.75
13.6
8.7
----
98.2
3.15
1.96
Good
Cake
3
9/1/70
32
1.0
...
6.3
2.75
8.7
118.8
3.88
2.14
Good
Cake
4
9/1/70
40
3.0
---
6.3
2.75
0
0
...
No
Cake
Formed
5
9/2/70
51.5
5.5
._.
6.3
2.40
21.3
10.0
110.0
3.22
2.30
Good
Cake
6
9/2/70
51.5
9.0
-__
6.3
2.40
21.3
0
...
No
Cake
Formed
7
9/9/70
19
2.0
23
6.3
2^38
21.6
19.1
110.0
4.80
2.86
Wet
Cake
8
9/9/70
36
5.5
12
6.3
2.38
21.3
10.0
116.0
3.88
1.96
Good
Cake
9
9/9/70
45
6.0
23
6.3
10.1
9.9
—
---
No
Cake
Formed
-------
TABLE 1 - APPENDIX F CONTD.
RUN NUMBER
DATE
Filter Flow Rate (gpm)
Drum Speed (ft/min.)
Chemical Addition
Lime (207. Solution) (ml/min)
FeCli (107. Solution) (ml/min)
pH of Feed After Chemical Addition
Feed Sludge
TSS , (mg/1)
VSS (ng-'l)
TDS (mg/1)
VDS (mg/1)
TS • TSS + TDS (mg/1)
Running Time (rain)
7. Solids in Cake
Filtrate
pH
TSS (n>sA)
VSS (mg/1)
IDS (mg/1)
VDS (mg/1)
TS - TSS + TDS (mg/1)
BOD (mg/1)
NHi - N (ng/1)
W>2- N (mg/1)
HOT - S fmi/n
10
9/9/70
1.6
1.49
136
3.9
23.504
16,202
342
132
23,846
30
18
6.2
116
89
2,200
796
2,316
17
...
—
11
9/15/70
1.6
1.49
130
130
11.6
20,480
15,032
536
76
21,016
30
24
11.4
279
147
3,378
1,016
3.657
15.5
...
1.0
12
9/17/70
1.6
1.49
190
no
11.7
15,004
11,116
540
48
15,544
23.75
22.5
11.4
283
175
602
80
885
11
0
1.0
13
9/24/70
1.6
1.49
65
65
10^6
16,648
8,480
708
508
17,356
20
17
9.5
98
54
1,208
352
1,306
15
— .
1.0
14
9/24/70
1.6
1.49
125
65
11.7
16,648
8,480
708
508
17.356
30
19.3
11.3
—
—
—
9.0
- —
0.5
15
9/24/70
1.6
1.49
185
65
11.9
16,648
8,480
708
508
17,356
20
22
11.5
...
...
....
16
9/25/70
1.6
1.49
135
135
11.7
18,136
13,504
680
240
18,816
30
24.5
11.3
—
11.0
...
-------
TABIZ I - APPENDIX F CONTD.
RUN NUMBER
DATE
TKN Cm^/11
TP fmf/n
ORTHO - f fmg/11
C^cle Tine ( min /rev )
T, 0>v wtj Solids to .Filter
Chemical Dosage (as % by wt. of dry
solids in sludge to filter)
Lime
Fed,
Polymer
Filtrated Collected (gal /hr)
Filter Cake Moisture
(Ib H,0/lb dry solids)
Cake Yield (Ib dry solids/hr/ft2)
Remarked
10
9/9/70
35
2.0
15
6.3
2.38
0
10.0
117.0
4.56
2.18
Good
Cake
11
9/15/70
35
13
6.3
2.10
24.2
11.4
106.8
3.17
1.50
Good
Cake
12
9/17/70
34
14
6.3
1.55
44.5
12.1
144.0
3.44
1.6B
Good
Cake
13
9/24/70
3.5
6.3
1.74
13.9
6.57
No
Cake
Formed
14
9/24/70
51.0
5.0
6.3
1.74
26.7
6.57
No
Cake
Formed
15
9/24/70
6.3
1.74
39.5
6.57
No
Cake
Formed
16
9/25/70
33.0
7.0
6.3
1.80
26.4
12.6
89.4
3.08
1.09
Thin
Cake
-------
TABLE I - APPENDIX F CONTD.
RUN NUMBER
DATE
FiUer Flow Rate (gpm)
Drum Speed (ft /tain. )
Chemical Addition
Lime (201 Solution) (ml/rain)
FeCL, (107. Solution) (ml/min)
Polymer
S-1938 (10% Solution) (ml/min.)
Calgon ST-287 (.027. Solution) (ml/min)
Calgon ST-287 (.01% Solution) (ml/min)
pH of Feed After Chemical Addition
Feed Sludge
TSS (mg/1)
vss (mg/i)
IDS (mg/1)
VDS (mg/1)
TS • TSS + TDS (mg/1)
Runnlnz Time (min)
TSoTias in Cike
Filtrate
pH
TSS (mg/1)
VSS (mg/1)
TDS (mg/1)
VDS (mg/1)
TS • TSS + IDS (mg/1)
BOD fme/11
NHi -K (mz/n
17
10/1/70
1.65
1.41
142
2.6
14,396
i n *?AA
680
248
15JD76
30
20.9
2.9
158
128
2,736
398
2,894
18
10/5/70
2.63
1.71
210
3.9
14,008
676
112
14,684
30
19.7
23
19
2,352
366
2,375
19
10/5/70
2.63
2.3
210
3.9
14,632
760
108
15^392
29
18.7
2.4
73
58
3,704
868
3,777
20
10/7/70
2.43
1.52
57
6.6
19,920
740
144
20,660
15
7.6
21
10/7/70
1.6
1.52
57
6.6
19,920
740
144
20,660
20
7.6
22
10/8/70
1.63
1.52
63
6.3
17,892
760
184
18.652
15
8.3
384
312
606
92
1A1
23
10/8/70
1.63
1.52
63
17^92
760
184
18.652
15
-------
TABLE I - APPENDIX F CONTD.
RUN NUMBER
DATE
NO, -NT (mg/1)
N03 -N (rag/I)
TKN (-8/D
IPT (na/11
ORTHO XP (ns/n
Cycle Time (air. -'rev)
% (by wt.) Solids to Filter
Chemical Dosage (as 7. by wt. of dry
solids in sludge to filter)
Lime
FeCU
Polymer
Filtrated Collected (eal /hr)
Filter Cake Moisture
Lib HsO/lb dry solids)
Cake Yield (Ib dry solids/hr/ft )
Remarks
17
10/1/70
....
----
-..-_
_,._-
6.65
1.51
0
15,9
- ^ - -
98.0
3.78
1.54
Good
Cake
18
10/5/70
....
5.5
1.47
0
15.2
114.6
4.08
1.71
Good
Cake
19
10/5/70
....
4.1
1.54
0
14.6
119.8
4.35
1.65
Good
Cake
20
10/7/70
....
6.18
2.07
2.85
No
Cake
Formed
21
10/7/70
....
6.18
2.07
2.85
No
Cake
Formed
22
10/8/70
6.18
1.86
0.0109
No
Cake
Formed
23
LO/8/70
6.18
1.86
.545
No
Cake
Formed
-------
TABLE I - APPENDIX F CONTD.
RUN NUMBER
DATE
Filter Flow Rate (Rpm)
Drum Speed (ft/mln)
Chemical Addition
Lime (201 Solution) (ml/min)
FeCl3 (10Z Solution) (ml/min)
Polymer
Dow A-23 (.057. Solution) (ml/min)
pB of Feed After Chemical Addition
Feed Sludge
TSS (niR/1)
VSS (mK/1)
IDS (mK/1)
YDS (mg/l)
TS - TSS + IDS (mg/l)
Running Time (rain)
% Solids in Cake
Filtrate
pH
TSS (mg/l)
VSS (mg/l)
TDS (mg/l)
VDS (mR/1)
TS - TSS + TDS (mg/1)
BOD (mg/l)
24
10/26/70
1.58
1.49
95
2.6
8,484
6,460
624
76
9,108
30
20.9
2.7
49
35
4,400
1,658
4.449
14
25
10/27/70
1.57
1.49
103
3,2
12,352
10,180
484
36
12,836
30
20
3.3
49
35
2.012
434
2,061
4
26
10/27/70
1.57
1.49
54
24
3.5
12,760
10,180
508
64
13,268
30
20.3
3.1
18
12
2,100
890
2,118
25
27
10/27/70
1.57
1.49
97
3.2
12,760
10,180
508
64
13,268
30
21.2
2.9
21
28
10/27/70
1.57
1.49
56
24
3.5
13,048
9,596
476
50
13,524
23.5
17.9
29
10/27/70
2.1
2^,76
110
3.10
14.508
10,732
368
14,876
30
18.8
182
1,954
722
2.136
-------
TABLE I - APPENDIX F CONTD.
RUN NUMBER
DATE
NHl - N fmj/ll
NO? - N Cmj/l"!
NOi - N Cm?/ 11
TKK (mg/1)
TP (mg/1)
ORTHO - P (mg/1)
Cycle Time (min/rev)
Z (by wt.) Solids to Filter
Chemical Dosage (as 7. by wt. of dry
solids in sludge to filter)
Lime
FeCl,
Polymer
Filtrated CDllected(gal /hf)
Filter CakeMoisture
(lb HpO/lb dry solids)
Cake Yield (Ib dry solids/hr/f t2)
Remarks
24
10/26/70
6.3
0.91
18.5
96.0
3.79
0.93
Thin
Cake
25
10/27/70
6.3
1.28
14.3
87.6
4.00
0.99
Thin
Cake
26
10/27/70
6.3
1.33
7.48
.0158
104.0
3.92
1.31
Good
Cake
27
10/27/70
_
6.3
1.33
13.0
97.2
3.72
1.36
Good
Cake
28
10/27/70
6.3
1.35
7.33
.0158
52.6
4.59
0.392
Very
Thin
Cake
29
10/27/70
3.4
1.49
9.82
101.0
4.31
2.12
Good
Cake
-------
TABLE I - APPENDIX F CONTD.
ars NUMBER
DATE
Filter Flow Rate (gpm)
Drum Speed (ft/min)
Chemical Addition
Lime (20% Solution) (ml/tnin)
Fed, (107. Solution)(tnl/min)
Polymer
Dow A-23 (.057. Solution)
(ral/siin)
pH of Feed After Chemical Add.
Feed Sludge
TSS (me/0
VSS (ms/1)
IDS (ra*/l)
VDS (BIR/l)
TS - TSS -t- IDS (mg/1)
Running Time (rain)
7. Solids in Cake
Filtrate
pH
ISS (mg/1)
VSS (mg/1)
IDS (mg/1)
VDS (mg/1)
TS - ISS + IDS (mg/1)
30
LO/28/70
1.58
1.49
100
3.5
16.168
11,788
768
120
16,936
30
21.2
3.4
34
44
2,556
846
880
31
10/28/70
1.58
1.49
51
24
16,168
11,788
768
120
16,936
20
32
11/2/70
2.00
2.5
91.6
4.2
11,780
L5.132
606
184
2,386
30
18.7
5.6
21
12
2,378
362
2,399
33
11/2/70
1.96
2.5
73.4
4.2
19,208
13,412
608
228
19,816
30
15.5
5.6
44
32
1,978
92
2,022
34
11/3/70
2.0
2.2
173
2.6
21,460
13,300
972
224
22,432
30
19.6
2.7
103
74
4,216
1,644
4,319
35
11/4/70
2.2
2.21
167
3.3
20,560
14,180
736
236
21,296
27.3
19.9
3.5
28
47
3,132
1,140
3,160
36
11/4/70
2.63
2.21
180
3.6
20,560
14,180
736
236
21,296
23
18
3.9
__ _ -
37
11/4/70
2.53
2.21
179
3.6
20,560
14,180
736
236
21,296
23
17.3
3.9
38
11/5/70
2,95
2.56
204
3.3
19,836
13,656
612
16
20,448
23
17.6
3.4
93
55
3,042
1,168
3.135
-------
TABLE I - APPENDIX F CONTD.
RUN NUMBER
DATE
BOD (mg/1)
NH, - N (mg/1)
N07- N (mg/1)
NOT N (mg/1)
TKN (mg/1)
TP (mg/11
ORTHO - P (mg/1)
Cycle Time (min/rev)
T. (by wt.) Solids to. Filter
Chemical Dosage (as % by wt.
of dry solids in sludge to
filter)
Lime
FeClj_^_
Folyner
Filtrated Collected(gal /hr)
Filter Cake Moisture
(Ib H20/lb dry
solids)
Cake Yield (Ib dry solids/hr/
ft*)
Remarks
30
0/28/70
22
—
—
...
6.3
1.69
10.5
92.0
3.71
i
1.54
Good
Cake
31
10/28/70
6.3
1.69
5.32
.0125
...
No
Cake
Formed
32
11/2/70
27
._.
...
3.77
2.24
5.77
111.6
4.35
2.62
Good
Cake
33
11/2/70
24
...
—
—
3.77
1.98
5.30
111.6
5.45
1.99
Wet
Cake
34
11/3/70
51
„
...
4.24
2.24
10.0
124.0
4.10
3.00
Good
Cake
35
11/4/70
15
...
4.24
2.13
10.0
136.0
4.02
2.67
Good
Cake
36
11/4/70
--.
4.24
2U3
9-.11
162.0
4.55
3.39
Good
Cake
37
11/4/70
...
...
4.24
2.13
9.4
162.0
4.78
3.36
Good
Cake
38
11/5/70
.
...
3.6f
2.04
9.6
162.0
4.6t
3.28
Good
Cake
-------
TABLE I - APPENDIX F CONTB.
RUN NUMBER
DATE
Filter Flow Rate (gpm)
Drum Speed (ft / min)
Chemical Addition
Lime (207. Solution) (ml/min)
FeCU(107. Solution) (ml/min}
pH of Feed After Chemical Addition
Peed Sludge
TSS (mg/1)
VSS (mg/1)
IDS (mg/1)
VDS (mg/1)
TS - TSS + IDS (mg/1)
Running Time (min)
T. Solids in Cake
Filtrate
pH
TSS (mg/1)
VSS (mg/1)
IDS (mg/1)
VDS (mg/1)
TS - TSS + TDS (mg/1)
BOD fa'1)
HHi- N (rng/1)
HOr N (mg/l)
HO}- N (mg/1)
39
11/10/70
3.97
4.0
300
3.6
19,974
14,728
604
80
20,578
16.5
15
64
47
3.222
1,072
3,286
...
40
11/10/70
3.97
4.0
300
3.6
19,974
14,728
604
80
20,578
30.0
14.9
64
47
3.222
1.072
3,286
41
11/10/70
300
3.6
19,974
14,728
604
80
20,578
15.0
14
64
47
3,222
1,072
3.286
__ T
11/11/70
4.0
4.0
260
3.1
16,464
12,212
572
152
17,036
16.5
16.6
3.0
31
17
2.860
876
2.891
34
—
iO
11/12/70
3.58
3.34
280
3.1
22,208
15,508
736
132
22,944
15.5
16
3.1
20
16
4,246
1,852
4,266
38
...
^4
11/30/70
2.5
2.5
150
155
11.8
10,372
6,636
706
78
,_11,078
20.0
20.5
11.4
515
258
512
168
1.027
—
—
-------
TAB1£ I - APPENDIX F CONTD.
RUN NUMBER
DATE
TKN (me/11
TP (me/1)
ORTHO - P (me/11
Cycle lime (min/rev)
% (bv wt.) Solids to Filter
Chemical Dosage (as 7, by wt. of
dry solids in sludge to filter)
Lime
Fed.
Polymer
Filtrated Collected (gal/hr)
Filter Cake Moisture
(Ib H,0/lb dry solids)
Cake Yield
( Ib dry solids / hr/' ft*)
Remarks
39
11/10/70
2.36
2.06
9.7
226.0
5.65
4.48
Thick
Wet
Cake
40
11/10/70
2.36
2.06
9.7
252.0
5.71
4.20
Thick
Wet
Cake
41
11/10/70
2.36
2.06
9.7
210.0
6.14
3.66
Thick
Wet
Cake
42
11/11/70
2.36
1.70
10.7
225.0
5.02
3.68
Thick
Wet
Cake
43
11/12/70
2.82
2.29
9.6
230.0
5.25
4.55
Thick
Wet
Cake
44
11/30/70
3.77
1.11
32.3
15.7
134.0
3.88
1.14
-------
TABLE 2 - APPENDIX F
THICKENED OXYGENATION SYSTEM WASTE ACTIVATED SLUDGE
RUN NUMBER
DATE
Filter Flow Rate (eonl
Drum Speed (ft/mirO
Chemical Addition
Lime (207. Solution) (ffll/n>in)
Fed, (107. Solution) rmi /min\
pH of Feed After Chemical Addition
Feed Sludge
TSS fmg/n
vss fmf/M
IDS fmf/11
VDS fmp/n
TS - TSS + TDS Cmg/n
Running Time (min)
Z Solids in Cake
Filtrate
pU
TSS (mg/1)
VSS (mg/1)
TDS (mg/1)
VDS (mg/1)
TS " TSS + TDS (mg/1)
1
11/23/70
2.55
2.4
400
3.1
39,066
26,508
1,540
404
40 , 606
20.4
12.6
3.1
102
91
5,214
1,866
5,316
2
11/23/70
2.2
.24
350
3.1
39,066
26,508
1,540
404
40,606
15.0
13
3.1
102
91
5,214
1,866
5,316
3
11/23/70
1.48
1.51
270
3.1
39,066
36,508
1,540
404
40,606
16.0
13.9
3.1
102
91
5,214
1,866
5,316
4
11/24/70
1.65
1.51
270
3.1
37.332
26,156
1,400
748
38,732
18.0
14.5
3.0
107
._
6,476
3,988
6,583
5
11/19/70
2.8
3.6
450
3.0
42,194
26,158
1,530
982
43 , 724
13.2
13.5
2.8
„
__
__
__
--
6
11/19/70
2.5
3.82
430
3.1
42,194
26,158
1,530
982
43,724
17.0
11-.^
70
74
12,036
3,468
12,106
-------
TABIi 2 - APPENDIX F CONTD.
RUN NUMBER
DATE
BOD (mg/1)
NH3 - N (mg/1)
NO, - N (mg/1)
NO, - N (mg/1)
TKN (mg/1)
TJ? (mg/1)
ORTHO -_P (mg/1)
Cycle Time. (min/rev)
% (by wt.) Solids to Filter
Chemical Dosage (as % by wt.
of dry solids in sludge to filter)
Lime
FeClj
Polymer
Filtrated Collected (eal/hrl
Filter Cake Moisture
Qb H-0/lb dry solidsl
Cake Yield
(Ib dry solids/hr/ft1)
Remarks
1
11/23/70
282
3.85
4.06
10.8
90.0
6.94
4.22
Wet Thick
Cake
2
11/23/70
282
3.85
4.06
11.0
100.0
6.69
3.96
Wet Thick
Cake
3
11/23/70
282
6.24
4J36
12.6
68.8
6.19
2.72
4
11/24/70
265
6.24
3.87
11.83
...
73.7
5.90
2.91
5
11/19/70
249
—
2.62
4.37
10.3
99.5
6.41
5.14
Wet Thick
Cake
6
11/19/70
—
2.47
4.37
11.0
78.9
6.05
5.12
Wet Thfck
Cake
-------
TABLE 3 - APPENDIX F
FILTER LEAF TEST DATA
(Filter Leaf Area = 0.1 ft2)
* Explained in back of Table 3-
Appendlx F
Date
10/ 13/70
10/13/70
10/13/70
10/13/70
LO/13/ 70
10/13/70
10/13 70
10/13/70
10/15/70
10/15/70
10/15/70
10/15/70
10/15/70
10/15/70
10/15/70
10/15/70
10/15/70
10/15/70
10/19/70
10/19/70
10/19/70
10/19/70
10/19/70
10/19/70
10/19/70
10/19'70
Run
No
1
2
3
4
5
6
7
8
1
2
3
4
5
6
7
8
9
10
1
2
3
4
5
6
7
8
?H of Chemically
Conuttioneu Raw wast
Activate,! S'uJge
2.9
2.9
2.9
2.9
3.5
3.5
4.0
4.2
3.0
3.2
3.3
3.0
6. 5
6.5
6.5
4.6
3.6
6.6
3.0
3.5
6.3
6.3
6.3
6.3
Chemical Addition
Type
'- 1*
A
A
A
B
C
D
A
A
A
A
A
A
E
F
F
G
H
F
I
H
H
H
E
J
K
L
Dosage
(mg/gof
drv solids
123
123
123
23.2
._
23.2
86
61
129
90
64
129
24.3
24.3
48.6
129
129
12.1
125
100
75
50
250
100
200
300
Type
2*
_
.
_
_
_
.
_
,
,
.
_
_
_
_
_
_
_
_
_
.
_
.
_
_
-
Dosage
(mg/gof
dry solid
_
_
_
_
.
_
_
_
_
_
_
_
_
_
_
_
.
_
_
.
.
_
.
_
-
Cake
1,
Solids
1
19.3
21.9
16.3
21.7
17.0
17.0
17.0
17.0
20.8
17.5
17.3
24.3
13.7
28.4
15.0
15.1
14.3
16.3
..
16.0
Yield ,
(Ib /hr/ ft')
0.243
0.278
0.433
0.0941
0.324
0.366
0.324
1.18
1.47
1.25
0.885
0.75
0.734
0.617
1.17
0.55
0.572
0.286
-.-
0.44
Filtrate
Volume
Collected (ml)
280
400
410
115
400
115
380
400
900
950
840
800
260
_
_
370
570
720
270
590
640
400
490
300
42
570
pH
3.3
3.2
3.4
6.2
5.5
5.0
4.2
5.5
2.9
3.1
3.5
3.5
7.0
_
_
5.9
5.5
7.0
3.3
3.9
5.4
5.9
6.6
7.0
7.0
6.9
Sludge
Type
ADS
ADS
ADS
ADS
ADS
ADS
ADS
ADS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
VAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
-------
TABLE 3 - APPENDIX F CONTD.
* Explained in back of Table 3
Appendix F
Date
10/19/70
10/19/70
10/20/70
10/20/70
10/20/70
10/20/70
10/20/70
10/20/70
10/20/70
10/20/70
10/21/70
10/21/70
10/21/70
10/2770
10/21/70
10/21/70
10/21/70
10/21/70
10/21/70
10/21/70
10/22/70
10/22/70
10/22/70
10/22/70
10/22/70
10/22/70
10/22'70
Run
No.
9
10
1
2
3
4
5
6
7
8
1
2
3
4
5
6
7
8
9
10
1
2
3
4
5
6
7
pH of Chemically
Conditioned Raw
Haste Activated Sludi
3.6
3.6
4.2
5.2
4.6
4.6
3.3
2.9
3.3
6.1
6.1
3.2
3.2
3.3
Chemical Addition
Type
1*
e
J
J
A
A
A
M
A
F
A
F
M
M
F
M
F
N
N
N
N
A
M
M
M
M
P
P
A
Dosage
(mg/g of
dry solids
100
100
109
s:
54
27
54
23
93
27
11
11
110
11
550
550
1100
275
275
82.5
8 ml
8 ml
8 ml
8 irl
Type
2*
A
A
_
_
_
A
M
A
_
A
A
F
A
A
A
A
.
A
A
,
A
F
F
A
A
A
-
Dosage
(mg/g of
drv solid
50
50
54
27*
47
__
109
55
330
55
L 55
55
55
._
82.5
82.5
__
20 ml
8 ml
8 ml '
20 ml.
Cake
7.
Solids
)
19.3
19.0
15.4
15.3
16.0
16.0
15.2
17.2
19.8
14.7
15.7
14.2
18.0
14.5
18.2
13.5
17.1
16.9
16.9
19.4
20.3
20.6
22.4
20.2
Yield
(Ib /hry ft )
0.638
0.528
0.722
0.622
0.526
0.514
0.652
0.722
1.81
0.717
1.08
0.60
0.64
0.69
0.915
0.383
1.43
1.71
1.58
1.50
1.43
1.25
1.32
1.36
Filtrate
Volume
Collected (ml)
490
450
430
540
590
610
525
570
775
625
680
260
600
580
590
590
580
830
790
820
690
350
230
660
590
775
620
PH
5.6
5.8
4.9
4.7
5.7
5.7
6.0
4.8
2.9
3.0
„
,
- ._
4.6
6.4
6.3
4.8
4.3
3.2
4.0
Sludge
Type
WAS
WAS
ADS
ADS
ADS
, ADS
ADS
WAS
WAS
ADS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
WAS
-------
TABLE 3 - APPENDIX F CONTD.
Legend for Chemical Addition Type 1 or 2
A - Fed
B - Aqua Floe
C - H2S04
D - H So + Aqua Floe
E - Rohm & Ha ss C-7
F - Dow C-31
G • Alum
I
J
K
L
M
N
P
• 107. Alum + H2SOA
- 4 ml Rohm & Hass + 21 ml H20 (17. C-7)
- 8 ml Rohm & Hass + 17 ml H20 (17. C-7)
- 12 ml Rohm & Hass + 13 ml HjO (17. C-7)
- A-23
» 20X-150
- Herk- 828.1
Sludge Type
ADS - AerobicaUy Digested Sludge
WAS - Waste Activated Sludge
-------
TABLE 4 - APPENDIX F
ANAEROBICALLY DIGESTED SLUDGE
RUN NUMBER
DATE
Filter Flow Rate (gpm)
Drum Speed (ft/min )
Chemical Addition
Lime (20% Solution) (ml/min)
FeCl, (107. Solution) (ml/min)
pH of Feed After Chemical Addition
Feed Sludge
TSS (mg/1)
VSS (mg/1)
TDS (m§/l)
YDS (mg/1)
TS = TSS + TDS (ms/1)
Running Time (min)
% Solids in Cake
Filtrate
PH
TSS (mg/1)
VSS (mg/1)
TDS (mg/1)
VDS (mg/1)
TS = TSS + TDS (mg/1)
BOD (mg/1)
NH - X (mg/1)
1
9/22/70
1.6
2.34
255
130
12
41,172
22,852
344
178
41.516
24.0
"20.3
11.7
294
27
3,806
238
4,100
_
362.5
2
9/21/70
1.6
1.49
133
138
7.7
41,172
22,852
344
178
41,516
20.0
7.7
339
218
2,622
1,384
2,961
.
125
3
9/21/70
1.6
1.49
265
275
8.8
41,172
22,852
344
178
41,516
20.0
6.8
_
_
-
.
-
.
114.5
4
9/28/70
3.05
3.9
900
280
12.2
44,208
11,968
736
236
44 , 944
20.0
20.0
12.0
830
135
4,376
1,816
5,206
.
-
-------
TABLE 4 - APPENDIX F CONTD.
RUN NUMBER
DATE
NO, - N me/1
NO 3 - N me/1
TKN me/1
TP me/1
ORTHO-P_ mg/1
Cycle Time (rain/rev)
% (by wt.) Solids to Kilter
Chemical Dosage (as % by wt. of dry
sol Ids in sludge to filter)
Lime
Fe CH
Polymer
Filtrated Collectedfgal /hr)
Filter Cake Moisture
(Ib H70 / Ib dry solids)
Cake Yield
(Ib dry solids/hr/ft2)
Remarks
1
9/22/70
_
1.0
_
1.5
.
4.03
4.15
22.7
5.42
__
148.7
3.88
5.56
Good
Cake
2
9/21/70
_
1.0
_
4.5
2.0
6.3
4.15
12.6
6.10
--
No Cake
Formed
3
9/21/70
_
1.0
_
5.0
2.0
6.3
4.15
25.2
12.2
--
No Cake
Formed
4
9/28/70
_
.
_
_
_
2.42
4.49
36.7
6.1
142.5
4.62
9.32
Wet Thick
Cake
-------
TABLE 5 - APPENDIX F
AEROBICALLY DIGESTED SLUDGE
RUN NUMBER
DATE
Filter Flow Rate (gpm)
Drum Speed ( ft/min y
Chemical Addition
Lime (20% Solution) (ml/min)
FeCl3 (107. Solution) („!/„,<„)
Polymer
Dow C-31 c.017. Solution) (ml/min)
Days in Digester
pH of Feed After Chemical Addition
Feed sludge
TSS (mg/1)
VSS (mg/1)
IDS (me/ 11
VDS (mg/1)
TS = TSS + IDS (me/ 11
_ Running Time (min)
% Solids in Cake
Filtrate
pH
TSS Cmg/11
VSS (mg/n
TDS Cme/ll
VDS (mg/1)
TS = TSS + TDS (mg/1)
1
9/29/70
1.68
1.49
138
130
6
9.2
18^420
7,224
1,652
952
20,072
30.0
17.2
9.0
868
880
3,932
1,568
4,800
7
9/29/70
1.68
1.49
138
130
6
8.8
18,420
7,224
1,652
952
20,072
L 30.0
18.1
8.6
144
103
4,822
4,028
4,966
3
10/21/70
1.59
1.52
100
26
7
4.0
30.0
18.5
4
10/21/70
1.59
1.52
50
10
7
25.0
--- -
5
10/21/70
1.59
1.32
100
7
3.2
_-- _
_._-_
_.
20.0
17.8
5.0
. - - -
_ . ~_
_ ...
6
11/9/70
2.5
2.56
300
8
2.5
31,124
0,536
1,692
740
2,816
20.0
... -
_
_ . _.
7
11/9/70
1.65
1.52
86
8
3.1
20.0
3.5
- . -_
-------
TABLE 5 - APPENDIX F CONTD.
; RUN NUMBER
DATE
BOD (mg/1)
NH} - N (BiR/U
NO, - N (m«/n
NO - N (mg/n
TKN (mg/1)
IP (mR/l)
ORTHO-P (mg/1)
Cycle Time, (min/rev)
Z (by wt.) Solids Co Filter
Chemical Dosage (as 7. by wt.
of dry solids in sludge to filter) '
[ Lime
FeCl.
Polymer
Filtrated Collected ( gal/hri
Filter CakeMoisture
(lb HnO/lb dry solids}
Cake yield
(lb dry solids/hr/ft2)
Remarks
1
9/29/70
....
....
_
6.3
2.0
17.1
7.5
57.6
4.81
0.583
Thin Cake
2
9/29/70
....
_ .
_ _ .
_. ..
6.3
2.0
17.1
7.5
60.0
4.52
0.533
Same
as
Previous
Run
3
10/21/70
., _
6.2
1.60
i-
11.1
0.273
62.0
4.40
0.862
Thin
Cake
4
10/21/70
....
6.2
1.60
5.0
0.110
No
Cake
Formed
5
10/21/70
_---
6.2
1.60
11.1
61.0
4.62
0.500
Thin
Cake
6
11/9/70
_ .
_
6.3
3.28
10.0
No
Cake
Formed
7
11/9/70
....
_-->
.
3.68
3.28
9.0
No
Cake
Formed
-------
TABLE 6 - APPENDIX F
NOMENCLATURE
A = Area of filtering surface (ft )
C. = Initial concentration of sludge fed to filter before conditioning
( % by wt. )
f\
g c= Newton's law conversion factor (ft-lb/lb force-sec )
K = Constant in Equation 10
k = Constant in Equation 9
L = Sludge cake yield ( lb/hr/ft2 )
m = Exponent on Ci in Equation 10
n = Exponent on t in Equation 10
2
p = Pressure drop through filter medium and sludge cake (lb force/ft )
r = Resistance of filter medium (ft )
•
S = Percent submergence of filter drum
t = Cycle time ( sec/rev or min/rev)
V = Volume of filtrate collected (ft3)
W = Weight of dry cake (lb)
3
w = Weight of dry cake per unit volume of filtrate^ lb /ft )
ol = Average specific cake resistance (Et/lb )
0 = Form time ( second or minuta)
^ = Viscosity of filtrate ( Ib/ft-sec )
155
-------
APPENDIX G - AEROBIC DIGESTION DATA
157
-------
TABLE 1 - APPENDIX G
SUMMARY OF' DIGESTER DATA
DAY
HUH NO
0
1
2
3
4
5
RUM NO.
0
1
2
3
4
5
um no.
0
1
2
3
4
5
6
TSS
. 1
24,812
24,292
20,840
20,100
18,012
17,868
, 2
-
14,518
12,692
12,276
12,416
11,576
3
-
15,996
18,440
13,760
16,828
-
14,492
VSS
17,632
17,420
16,188
13,608
11,672
7,056
-
10,660
-
8,800
9,308
8,048
-
14,656
13,768
12,992
12,208
12,956
10,204
IDS
880
692
764
1072
1436
L548
-
668
808
856
1020
940
-
700
736
824
1116
-
956
VDS
236
240
-
496
780
820
-
92
220
284
476
308
-
160
180
140
472
-
304
D.O. 02 UPTAKE CODg,
(All In mg/1)
7.5
10.5 68.7 142
37. 196
15.
14.5 - 319
15.0 - 173
-
39.2
18.5
25.8 204
-
119
.
6.0 68.7 93
4.0 51.0 145
3.3 62.5 209
4.5 15.2 343
9.5 - 154
14.1 93
TPS* (NH,-^ (N02-N)S, (N03-N),
...
17.4 26.8 0.0 .1
23.1 - 0.1 11.1
-
42.5 268. 2.0 150.
43.9 275. 5.5 172.
.
-
-
4.4 13.1 0.0 .2
-
32.8 12.8 1.6 47.6
-
29.0 35.4 0.0 0.4
26.9 61.5 0.1 6.0
39. 138.4 0.1 14.2
50.7 158.5 0.1 46.2
-
44.6 178.5 0.3 52.6
5* TIOS
.
32.4
109.7
-
257.
260.
.
-
-
16.9
-
114.7
-
31.7
72.3
122.2
186.9
-
207.8
* Soluble
-------
PAY
TSS
vss
IDS
TABLE 1 - APPENDIX G CONTD.
D.O. 02 I'PTAKE COD
RUN
0
I
2
3
4
5
6
7
RUN
0
I
2
3
4
5
6
7
NO. 4
-
28,348
25,448
23,428
24,024
-
20,464
20,472
NO. 5
-
21,576
17,816
15,464
16,892
14,976
17,120
15,120
-
20,840.
18,568
16,588
16,972
15,004
14,26V"
14,184
-
15,696
-
11,924
12,276
10,984
14,976
11,192
-
600
592
1156
-
1384
2564
2512
-
680
664
868
1236
1732 -
1356
1624
-
224
56
404
-
792
1248
1276
-
424
140
448
464
828
472
1188
2.0
3.5
5.5
10.0
15
15
15
-
4.0
3.5
4.0
-
-
8.7
8.5
7.7
-
66.7
-
46.7
-
10.0
23.4
-
142
133.2
76.0
-
29.8
32.4
25.2
18.0.
-
178
214
231
207
196
196
268
-
429
-
298
-
-
326
265
-
29
47.1
44
61.0
75.9
88.8
103.2
-
46.5
-
49.5
-
-
66.
72.
-
59.8
39.5
101.9
-
145.7
183.2
204.1
-
95.4
-
145.8
-
-
169
162.
-
0.1
0.0
0.1
-
0.0
0.0
0.0
-
0.0
-
.1
-
-
7.2
7.9
-
0.2
0.3
39.4
-
225.7
293.2
301.8
-
0.0
-
14.9
-
-
91.
105.
-
57.3
92.6
92.6
-
167.5
186.0
202.7
-
146.3
-
178.2
-
-
185.
168.
* Soluble
-------
TABLE I - APPENDIX G CONTD.
DAY-
RUN
0
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
22
23
24
TSS
NO. 6
19,484
17,432
-
-
15,328
14,620
13,880
13,772
13,600
-
13,540
12,916
12,584
12,372
11,964
12,176
13,200
13,236
11,280
-
11,124
11,336
-
-
vss
14,436
13,028
-
-
11,232
10,769
10,036
9,860
10,084
-
9,720
8,976
9,076
9,316
8,588
8,852
9,332
13,124
8,016
8,336
8,452
8,220
-
IDS
-
732
-
-
-
1243
1492
1396
1416
1236
-
2508
1968
2108
1712
2732
3072
1784
5004
2724
2704
2976
2444
-
-
VDS
-
304
-
-
-
124
872
1028
704
424
-
1456
1052
1280
952
1896
2248
828
2976
1936
1804
2028
996
-
-
D.O.
-
7.5
7.5
8.7
8.0
7.5
8.5
7.2
7.0
7.6
7.0
6.7
-
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
°2 ^PTAKE
-
82.0
-
53.8
38.5
28.8
21.6
35.0
14.1
9.6
13.9
8.3
11.3
-
-
-
8.7
16.0
-
-
4.1
-
-
-
-
00 V
-
226.
255
-
313
-
309
300
309
318
-
-
382
464
418
464
563
563
481
527
546
685
611
518
-
T ps*
-
23.
40
-
10.2
-
7.5
7.9
8.3
8.5
-
-
93.8
98.4
99.6
100.2
105.7
105.7
82.7
95.6
96.7
109.2
98.3
95.3
-
(NH,-N)S,
-
21
63
-
17.4
-
19.3
21.2
24.1
25.4
-
-
291.
293.7
305.
301.2
312.8
310.9
296.
306.8
313.8
306.8
302.2
298.
.
„ (NOj-N)
-
0.0
0.0
-
.2
-
7.1
1.5
4.9
4.3
-
-
11.7
13.4
15.1
16.5
18.5
19.6
-
20.8
21.1
21.8
22.2
23.1
_
s* W>-N>s
-
0.0
17.
-
18.5
-
14.3
20.5
20.2
21.1
-
-
294.3
295.9
308.2
304.9
314.4
303.7
-
322.6
320.
321.7
314.4
311.2
_.
!* TkNS«
-
52
67
-
18.0
-
18.8
23.5
22.7
22.5
-
-
285.9
280.3
290.8
301.7
303.3
293.2
254.9
258.9
270.7
258.
248.
256.4
.
*Saluble
-------
TABLE 1 - APPENDIX G CONTD.
D-°- 0: '-HTAKE CODS> TP^ (NH3-N)S» (NO,-N)S^ (N03-N)S* TK:-g^
RL'X
0
1
2
3
4
5
6
7
S
RUN
0
1
2
3
4
5
6
7
8
9
10
NO. 7
43,672
42,016
39,548
36,716
27,736
34,736
33,660
32,820
31,124
NO. 8
30,752
-
29,620
-
-
-
24,864
24,372
24,368
22,946
22,524
28,980
28,625
28,012
25,696
24,772
23,844
22,512
21,580
20,536
20,780
-
20,592
-
-
-
13,496
16,980
16,664
15,960
15,796
1232
680
428
500
1860
1688
1764
3872
1692
1580
1560
1512
1628
-
-
1642
1712
1814
1720
1804
380
100
144
232
996
776
724
510
740
792
848
660
972
-
-
782
1168
818
636
1040
-
-
2.8
3.0
7.0
9.0
8.5
11.0
-
-
3.5
6.0
8.5
10.5
15
15
15
14
-
-
-
750
842
842.
-
23.4 705
27.6 610
676
611
611
-
-
-
39.0
25.7 530
418
484
474
557
680
.
55.4
73.
53.1
57.
55.6
59.7
-
-
49.8
59.9
86.8
-
-
63.3
77.2
79.1
94.3
88.8
116.2
. .
73.4 .2 .4
87.6 .3 .7
140.7 .2 .8
182.1 .2 1.1
237. .2 9.6
L37.4
-
- -
-
-
-
.
.
232.2 0.1 7.9
255.6 0.2 10.4
272.4 0.8 12.4
279.5 3.7 13.8
274.8 3.2 15.9
339.7 24.5 7.4
.
168.
275.5
255.5
290.8
289.2
311.7
-
-
-
-
-
-
-
266.2
278.3
293.5
285.8
287.5
347.9
*Soluble
-------
TABLE 1 - APPENDIX G CONTD.
VSS IDS \™ D.O. 02 I'PTAKE COD, T
RUN
11
12
13
14
15
16
17
18
NO. 8 CONTD.
21,916
-
-
20,396
21,056
20,144
20,400
-
15,224
-
-
14,848
15,060
13,628
13,668
-
3476
-
1400
2424
1492
2408
1126
1030
1612
-
920
1552
764
1320
482
458
-
15
15
15
14.8
14.0
14.0
14.8
-
-
21.4 540
548
400
453
514
480
-
-
73.4
86.1
63.8
83.7
89.2
89.2
-
-
244.7
259.3
274.2
259.9
268.3
268.9
-
-
0.3
0.6
0.5
0.1
0.2
0.1
-
-
22.5
26.0
27.2
26.5
28.6
29.1
-
-
261.6
239.
244.9
265.2
265.2
278.9
Soluble
-------
TABLE 2 - APPENDIX G
DIGESTER RUN NO. 6
Time of Aeration
(days)
0
2
5
6
7
8
9
11
12
13
14
15
16
17
20
21
22
VSS (mg/1)
13,200
13,028
11,232
10,769
10,036
9,860
10,084
9,720
8,976
9,076
9,316
8,588
8,852
9,332
8,336
8,452
8,220
Degradable
VSS (mg/1)
4980
4808
3012
2549
1816
1640
1864
1500
756
856
1096
368
632
1112
116
232
0
Degradable
(vss)t=t
Degradable
t=o
1.0
0.965
0.6048
0.512
0.365
0.329
0.374
0.301
0.152
0.172
0.220
0.074
0.127
0.223
0.0233
0.0465
0
163
-------
TABLE 3 - APPENDIX G
DIGESTION RUN NO. 8
Time of Aeration
( days)
0
2
7
8
9
10
11
14
15
17
VSS (mg/1)
20,780
20,592
16,980
16,664
15,960
15,796
15,224
14,848
15,060
13,668
Degradahle
VSS (mg/1)
7112
6924
3312
2976
2292
2128
1556
1180
1392
0
Degradable
(vss>t=t
Degradable
i.o
0 .974
0-467
0-418
0.322
0 .299
0 .2.9
0 .166
0.196
0
164
-------
1
Accession Number
w
2 1 Subject Field & Group
O5G
SELECTED WATER RESOURCES ABSTRACTS
INPUT TRANSACTION FORM
5 .
Union Carbide Corporation, Tonawanda, New York
Linde Division
Title
10 \
Continued Evaluation of Oxygen Use in Conventional
Activated Sludge Processing
M. A. McDowell
N. P. Vahldieck
E. A. Wilcox
K. W. Young
f£ I
- 1
Project Designation
Project No. 17050 DNW - Contract No. 14-12-867
Note
22
Citation
23
Descriptors (Starred First)
*Waste Treatment, *Activated Sludge, *Oxygen, Secondary Treatment,
Municipal Wastewater, Aerobic Digestion, Vacuum Filtration, Sludge Handling,
Sludge Disposal.
OC Identifiers (Starred First)
*High-Purity Oxygen, *Batavia, New York, Union Carbide Oxygenation System
07 Abstract
A process for treating municipal wastewater using high-purity oxygen in the
activated sludge process was further evaluated at Batavia, New York in a full scale
wastewater treatment plant. The present evaluation places emphasis on the production
rate, filtration, and digestion characteristics of oxygenated waste activated sludge.
Operation was conducted at two different system loadings over a period of three months.
Sludge production data obtained during each period of operation was verified by perform-
ing solids and mass balances over the entire system on a continuous basis. The excess
solids production data obtained in this contract confirm the observations made during
the initial contract that an oxygenation system produces up to 50% less solids than a
conventional air aeration system.
Waste sludge from the oxygenation system clarifier underflow was filtered directly on
a pilot-scale vacuum filter. A filter cake was produced which disengaged easily, and
had a solids content of 15-24 percent. Filter yields as high as 4. 5 pounds per hour
per square foot of filter area were obtained.
Waste activated sludge was aerobically digested using high-purity oxygen.
percent reduction in volatile suspended solids was obtained in 7-9 days.
A 20-30
Abstractor
E.
A.
Wilcox
'"If
i ti/tion
nion
Carbide
Corporation
- Linde
Divis
ion
WR 102 (REV, JULY 19*9)
WRSJ C
SEND WITH COPY OF DOCUMENT, TO: WATER RESOURCES SCIENTIFIC INFOKMATION C E f
U.S. DEPARTMENT OF THE INTERIOR
WASHINGTON, D. C. 20240
U.S. GOVERNMENT PRINTING OFFICE: 1972-484-484/176
------- |