WATER POLLUTION CONTROL RESEARCH SERIES • 17010 ECZ 02/71
Wastewater Ammonia Removal
by Ion Exchange
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 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, Research Information Division, Research and
Monitoring, Environmental Protection Agency, Washington,
D. C. 20460.
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WASTEWATER AMMONIA REMOVAL BY ION EXCHANGE
Mobile Pilot Plant Studies and Process Design
with Electrochemical Renovation of Regenerant
by
Battelle-Northwest
Richland, Washington 99352
Project #17010 ECZ
Contract #14-12-579
Process Design with Air Stripping
Renovation of Regenerant
by
Soutr. Tahoe Public Utility District
South Lake Tahoe, California 95705
Project #17010 EEZ
Contract #14-12-561
for the
ENVIRONMENTAL PROTECTION AGENCY
February 1971
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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 neces-
sarily reflect the views and policies of the
Environmental Protection Agency, nor does mention
of trade names or commercial products constitute
endorsement or recommendation for use.
11
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ABSTRACT
Pilot plant investigations were conducted on the ion exchange removal of
ammonia-nitrogen from clarified and carbon-treated secondary effluents
and from clarified raw sewage. The ion exchange process utilized clinop-
tilolite, a natural zeolite. Average ammonia removals from low magnesium
wastewaters were in the range of 93% to 97%. With a wastewater Mg concen-
tration of 20 mg/1, solids formation presented problems but they appear
surmountable. The primary method used for regenerant renovation was air
stripping with which a 2N regenerant at a pH of 11.5 is recommended. Elec-
trolytic regenerant renovation using a neutral solution that is less prone
to solids formation was also piloted during the project.
Two process designs are included giving cost estimates for ion exchange
ammonia removal from tertiary effluent. With capital costs amortized at
6% for 20 years, the total cost to remove ammonia from 1000 gal. of tertiary
effluent is 14-.8?! for a 7-5 m§cL plant using regenerant air stripping and 12,70
for a 10 mgd plant using electrolytic regenerant renovation. The 7-5 m§d. de-
sign was prepared by South Tahoe Public Utility District under EPA Project
Number 17010 EEZ and is included for convenience. Other work discussed in
the report was performed by Battelle-Northwest under EPA Project Number
17010 ECZ.
This report was submitted in fulfillment of Project Number 17010 ECZ,
Contract 1^4-12-579> under the sponsorship of the Water Quality Office,
Environmental Protection Agency.
111
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CONTENTS
SUMMARY AND CONCLUSIONS
RECOMMENDATIONS
INTRODUCTION
PILOT PLANT OPERATION
ION EXCHANGE EQUILIBRIA
COLUMN OPERATION
Single Column Loading Studies
Two Column Semi-Countercurrent Operation
Zeolite Attrition
REGENERATION STUDIES
Normal Pilot Plant Regeneration
Pilot Plant Batch Recycle Regeneration
REGENERANT RECOVERY
Air Stripping of Regenerant
Electrochemical Renovation of Regenerant
ACKNOWLEDGMENTS
REFERENCES
APPENDIX A
APPENDIX B -
APPENDIX C -
Sample Calculation of Ammonium Ion Loading
Using Activity Coefficients
Preliminary Design of a 10 mgd Ammonia
Removal Plant Utilizing Electrolytic
Renovation of Spent Regenerant
Engineering Design of a 7 = 5 mgd Ammonia
Removal Plant Utilizing Air Stripping for
Recovery of Spent Regenerant
Page
1
3
5
7
11
15
15
19
2^
27
27
32
39
39
39
50
55
62
V
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FIGURES
No. jfog
1 FLOWSHEET FOR AMMONIA SELECTIVE ION EXCHANGE PROCESS 8
2 PHOTOGRAPH OF MOBILE PILOT PLANT IN OPERATION AT SOUTH 9
LAKE TAHOE
3 SELECTIVITY COEFFICIENTS VS. CONCENTRATION RATIOS OF 12
SODIUM OR POTASSIUM AND AMMONIUM IN THE EQUILIBRIUM
SOLUTION WITH HECTOR CLINOPTILOLITE AT 23°C FOR THE
REACTION (Y) 4- (NH. ) = (NH. ) + Y
Z 4 N 4 Z N
4 SELECTIVITY COEFFICIENTS VS. CONCENTRATION RATIOS OF 13
CALCIUM OR MAGNESIUM AND AMMONIUM IN THE EQUILIBRIUM
SOLUTION WITH HECTOR CLINOPTILOLITE AT 23°C FOR THE
REACTION (X)z + 2(NH4)N = 2(NH4)Z + XN
5 MINIMUM BED VOLUMES AS A FUNCTION OF NH3-N CONCENTRATION 17
6 AMMONIA BREAKTHROUGH CURVES FOR A 6 FT CLINOPTILOLITE BED 18
AT VARIOUS FLOW RATES
7 EFFECT OF BED DEPTH ON AMMONIA BREAKTHROUGH AT 4 . 8 gpm/ft2 20
FLOW RATE
8 EFFECT OF BED DEPTH ON AMMONIA BREAKTHROUGH AT 9 . 7 bv/hr 21
9 AMMONIA BREAKTHROUGH CURVES FOR 1ST AND 2ND COLUMNS IN SERIES 23
WITH TAHOE TERTIARY EFFLUENT
10 PHOTOGRAPH OF WHITE MATERIAL ON TOP OF A BED OF CLINOPTILOLITE 28
AT POMONA
11 EFFECT OF PURE WATER BACKWASH RATE ON SUBSEQUENT NH3~N BREAK- 29
THROUGH
12 EFFECT OF ACID WASH ON SUBSEQUENT NH3-N BREAKTHROUGH 30
13 AMMONIA BREAKTHROUGH CURVES FOR COLUMNS FOLLOWING REGENERATION 31
WITH UNCLARIFIED (1) AND CLARIFIED (2) REGENERANT WITH HIGH
MAGNESIUM CONTENT
14 FIRST BATCH RECYCLE ELUTION CURVES 33
15 FIRST AND SECOND BATCH RECYCLE ELUTION CURVES 34
16 REGENERANT BATCH RECYCLE WITH 3.6 BED VOLUMES OF 1 M CaCl2-0.2 M 36
NaCl AT pH 11
VI
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D .
Page
7 EFFECT OF REGENERANT VOLUME AND pH ON AMMONIA ELUTION 37
8 EFFECT OF Mg+2 ON CELL RESISTANCE 41
9 EFFECT OF H+ ACTIVITY AND FLOW RATE ON CELL RESISTANCE 42
0 BREAKPOINT CHLORINATION OF REGENERANT BATCH //I, RUN 3 45
vii
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TABLES
No. Page
1 Typical Composition of Feed Streams Ic
2 Tahoe Performance Data for Runs with Two Columns in Sefies 22
3 Performance Data for Seventeen Runs at Blue Plains with 25
Two Columns in Series
Vlll
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SUMMARY AND CONCLUSIONS
1. The selective ion exchange process for ammonia nitrogen (commonly
called ammonia) removal from wastewater is believed to be best suited
for use in those areas which experience prolonged periods of freezing
weather during winter and where very high degrees of removal must be
consistently maintained. Other processes such as air stripping^!/
alone and biological nitrification-denitrification(2) may be used in
warm climates at a lower cost but at somewhat lower efficiency.
2. Process designs provided costs for two alternate ion exchange methods
of ammonia removal from tertiary effluent. From a Battelle design
utilizing electrolytic regenerant renovation (this project), the cost
per thousand gallons was 9.04
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5. Batch regeneration by recirculating 4 bed volumes of 1 M GaCl2, 0.2M
NaCl and pH 11.5 (adjusted with lime) solution appears to minimize
regenerant volume. However, a wide range of concentrations was not
investigated and steady-state conditions may not have been reached. The
regenerant was chosen to include: (1) significant Na which has been
found to lengthen the service cycle and shorten the.elution cycle; (2) 1M
(2N) Ca concentration to provide for high elution capacity and approximate
what would be expected with continued lime pH adjustment in service; and
(3) high pH to shift the equilibrium toward NH3 production to facilitate
air stripping. Further optimization of the regenerant composition may
result from experience. Wasting of regenerant may be necessary if
undesirable build-up in concentration occurs. On the other hand, con-
siderable salt addition may be necessary due to dilution of the regenerant
during use. Allowances for NaCl and lime additions used in the designs
discussed in this report may have to be adjusted.
' ' +2
6. Processing of wastewaters with high Mg concentrations may require
clarification of the regenerant to avoid plugging the bed with Mg(OH)2-
Additional work is needed to verify this approach.
7. Processing of clarified and filtered raw sewage appears to cause
some biological growth in the zeolite beds but is adequately removed
during the regeneration cycle.
8. Zeolite attrition was 0.17% per cycle when using high pH regeneration
due to backwashing required for solids removal. Attrition would be very
low for the neutral regeneration proposed with electrochemical regenerant
renovation. Under neutral conditions much less solids would form and
high rate backwashing would be unnecessary. Zeolite ammonia capacity
does not change significantly with service.
9. Pilot studies demonstrated that air stripping of regenerant is
practical. The calcium carbonate scale formed on the stripping column
packing did not interfere with the stripping efficiency for operating
periods up to 65 days. The physical character of the scale varied
from flaky to hard. Water spraying removed the flaky scale to a large
degree, but water fluidization of the packing was required for removal
of the hard scale.
10. Laboratory studies and preliminary pilot studies demonstrated the
feasibility of electrochemical renovation of regenerant. The electrical
energy required to remove one gram of NH3-N varied from 35-54 watt hours.
White scale formation occurred on the cell cathode. This is expected to
be minimized, however, by promoting turbulence in the cell with baffles
or other cell modifications.
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RECOMMENDATIONS
Further development of the electrochemical method of removing the ammonia
from the spent regenerant is recommended. The major advantages of this
approach relative to that of air stripping of the spent regenerant are:
1. No precipitation occurs in the zeolite bed during regeneration
because neutral solutions are used,
2. No atmospheric disposal of ammonia is necessary,
3. Overall scaling problems associated with the use of lime
are eliminated.
Preliminary cost studies indicate that electrochemical renovation of spent
regenerant ,will be competitive with air stripping renovation.
The development program recommended for electrochemical treatment of spent
regenerant from the selective ion exchange process should include studies
to determine electrode life, and optimum current density, pH, salt concen-
tration and temperature. Methods to control calcium and magnesium hydroxide
scaling on the cathode should be evaluated to minimize cell resistance.
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INTRODUCTION
The results of laboratory and initial pilot plant studies on ammonia
removal by selective ion exchange have been reported previously(3»^)-
A naturally occurring zeolite, clinoptilolite, is employed as the ion
exchange medium which preferentially sorbs ammonium ions in the presence
of sodium, calcium, and magnesium ions. The zeolite can be regenerated
with solutions containing high concentrations of calcium ions. Spent
regenerant solutions can be renovated for reuse by air stripping or
electrolysis to remove ammonia. Basically, the selective ion exchange
process concentrates the ammonia into a relatively small volume of
liquid (regenerant) which can be: (1) air stripped even in cold weather
using low heat input to prevent freezing and to maintain high ammonia
removal efficiency; or (2) electrolyzed to convert the ammonia into
innocuous nitrogen gas.
Consistently high ammonia removals from clarified trickling filter
effluent were previously demonstrated in the laboratory^). The ammonia
removals varied from about 95% for a single zeolite bed with an output
of 150 bed volumes to more than 99% removal for two beds in semi-countercurrent
series operation with an output of 200 to 300 bed volumes.
The main objective of this program is demonstration of the use of the
selective ion exchange process on an engineering scale for removing
ammonia from a variety of wastewaters. A 100,000 gpd mobile pilot plant
was employed in this effort which included operations at the South Tahoe
Public Utility wastewater treatment plant at South Lake Tahoe, California;
the Pomona Wastewater Treatment Plant at Pomona, California; and the Joint
WQO-DC Pilot Plant at Blue Plains in Washington, D. C. Wastewaters
encountered in the pilot plant studies include clarified (Pomona) and
carbon treated (Tahoe and Pomona) activated sludge plant effluents and
clarified raw sewage (Blue Plains).
The above demonstration sites were selected to give a wide range in
dissolved organic concentrations and dissolved salts concentrations,
particularly magnesium salts. The dissolved organics may foul the zeolites
directly or indirectly by supporting biological growths which cover the
zeolite particles. Magnesium precipitates in the zeolite bed during
regeneration with lime and may cause operational problems where the
concentration is high.
Two process designs have been developed for zeolite ammonia removal from
tertiary effluent. During the present Battelle project, a 10 mgd plant
utilizing electrochemical renovation of the regenerant was designed. This
design paralleled one developed earlier for a 7.5 mgd plant using regenerant
air stripping by the South Tahoe Public Utility District under WQO/EPA
Contract No. 14-12-561. The 10 mgd Battelle design is Appendix B in this
present report and the South Tahoe project report is included as Appendix C.
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PILOT PLANT OPERATION
O)
The mobile pilot plant used in these studies has "been described previously
The water treatment unit (Recla-Mate SWB tertiary sewage treatment plant
manufactured by Neptune MicroFloc, Incorporated, of Corvallis, Oregon) was
not used for the most part because clarified water was available at the
demonstration sites. The 39 in- diameter ion exchange vessels were converted
from 750 gallons capacity to 500 gallons capacity by leaving off the top sec-
tions. This reduction in size greatly facilitated assembly and disassembly of
the ion exchange vessels by eliminating the use of a crane to lift the top
section to and from the roof of the trailer. The 500 gallon capacity was en-
tirely satisfactory for the demonstration program. A flow diagram for the
selective ion exchange process is illustrated in Figure 1 and a photograph of
the mobile pilot plant in operation at South Lake Tahoe is given in Figure 2.
(R)
The stripping tower packing consisted of 1 in. polypropylene Intalox^
saddles for ease in removal and repacking. Column diameter was ^3 in. and
packing depth was 7 ft.
Electrochemical renovation of regenerant involved the use of two 500 amp
electrolysis cells (manufactured by Pacific Engineering and Production
Company of Henderson, Nevada). The electrolysis cell consists of a lead
dioxide coated graphite anode, ^ l/^ in. diameter by ij-5 in- l°ng» placed in
a copper can that serves as the cathode.
The feed streams were pumped downflow through either a single zeolite bed
or two beds in series. Regenerant solutions were pumped upflow through
the beds at rates sufficient to remove precipitated or filtered solids
from the beds. During semi-countercurrent operation, which permits greater
utilization of the available ion exchange capacity, the loaded zeolite bed
(first in series) was removed from service for regeneration while a freshly
regenerated bed was placed at the end of the series.
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FILTRATION AND ION EXCHANGE
Wastewater
Flocculation, Sedimentation
and Mixed Media Filtration
Filter Pump
Backwash
<«„-.
Filtered
Water Storage
n
Filter
Backwash Pump
CD
cn
•4-
o>
en
tc
c.
o
["Air and A
m
I
monia Exhaust
-J
Clean Water
for Reuse I
Stripping
Tower
Tn Tm T*
jr^-J
Cm
Liqui
Sa
"I
1
t
L
P l-ino
Main Ion
Exchange Pump
Regenerant
and Ammonia
Lime,
Salt Addition
Air In
Air Blower
Make-up Tank
ZEOLITE REGENERATION PROCESS
Regeneration
Pump
FIGURE 1. FLOWSHEET FOR AMMONIA SELECTIVE ION EXCHANGE PROCESS
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FIGURE 2. PHOTOGRAPH OF MOBILE PILOT PLANT IN
OPERATION AT SOUTH LAKE TAHOE
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ION EXCHANGE EQUILIBRIA
The ion exchange equilibria for the systems NH4-Na , NH^-K , NH4-Ca
and NH^-Mg+2, with clinoptilolite and other zeolites, have been reported
by Ames'3'. The original data reported by Ames has been extended to
include higher Ca+2: NH^ and Na+: NH^ ratios. Plots of the NH^ selectivity
coefficients (defined in Appendix A) vs. the solution concentration ratios
of the cations are shown in Figures 3 and 4. This additional data is
useful for computing the equilbria in regenerant solutions. However,
standard solutions of 0.1N_ were used in obtaining the data. Corrections
for activity differences were needed to improve accuracy when using the
data in Figures 3 and 4 to predict maximum wastewater ammonia loadings
on clinoptilolite. An example of the computation is given in Appendix A.
11
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103
102
10
Y- Na
0.01
0.1
10
100
1000
(YN)
P4JJ)
FIGURE 3. SELECTIVITY COEFFICIENTS VS. CONCENTRATION RATIOS OF SODIUM OR POTASSIUM
AND AMMONIUM IN THE EQUILIBRIUM SOLUTION WITH HECTOR CLINOPTILOLITE AT
23°C FOR THE REACTION (Y) + (NH ) = (NH . ) + Y
(NH )
T IN
= (NH . )
T" Lj
IN
-------
105
Iff4
103
102
10
10
102 103 104 105 106 107
(NH4)
FIGURE 4. SELECTIVITY COEFFICIENTS VS. CONCENTRATION RATIOS OF CALCIUM OR MAGNESIUM
AND AMMONIUM IN THE EQUILIBRIUM SOLUTION WITH HECTOR CLINOPTILOLITE AT
23°C FOR THE REACTION (X) + 2(NH ) =
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COLUMN OPERATION
Single Column Loading Studies
Column loading studies were carried out to establish the volume of feed
water that can be processed through a zeolite bed until significant ammonia
breakthrough occurs. The typical composition of the various feed streams
employed in these studies are listed in Table 1.
TABLE 1
TYPICAL COMPOSITION OF FEED STREAMS
Activated Sludge Plant Effluent
Tahoe
Carbon Treated Pomona*
NHo-N mg/liter
Na mg/liter
K mg/liter
Mg mg/liter
Ca mg/liter
15
^
10
1
51
16
120
18
20
Clarified Raw Sewage
Blue Plains
12
35
9
0.2
30
pH
Range
Avg.
7-
6.5-8.2*
6.9, 7.8*
7-9
7-9
COD
TDS
11
325
10
700
50
250
*NOTE: Approximately half of the runs at Pomona were made with carbon treated
secondary and the others with alum coagulated secondary. The average
pH of the carbon column effluent was 7-8 and the average pH of the alum
coagulated secondary was 6.9-
15
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The equilibrium NHo-N bed loading computed for each of the wastewaters
listed in Table 1 is ^.1 g/1, 3-9 g/1. and 4-. 3 g/1. respectively, for
Tahoe, Pomona, and Blue Plains. Figure 5 presents equilibrium bed load-
ing in an alternate way. The minimum bed volumes required to attain
equilibrium NH^-N loading are expressed as a function of the NH-j-N con-
centration in the wastewater with the concentrations of metal ions held
constant. For actual operation, the bed volume values given in Figure 5
will normally represent the 50 percent breakthrough point where the ef-
fluent concentration is 50% of the feed concentration.
Tahoe and Blue Plains column loading data are given below. However,
Pomona loading data are thought to be atypical because of magnesium hy-
droxide formation during' regeneration, and are not reported here. Ways
of dealing with the magnesium hydroxide solids were investigated at
Pomona but time did not permit obtaining loading data under realistic opera-
ting conditions. The Pomona work is discussed in the section on Regenera-
tion Studies.
Ammonia breakthrough curves for a single 6 ft deep bed of clinoptilolite
are illustrated in Figure 6 for Tahoe tertiary effluent with flow rates
varying from 6.5 to 9-7 bv/hr (bed volumes per hour) with 15 to 17 mg/1
NH3~N in the feed stream. A throughput value of 150 bed volumes is
recommended for design. The average NH^-N concentration of the total
effluent to that point would be about 1 mg/1 or less. Follow up break-
point chlorination would probably be more effective for removing the
residual, if required, than greater ion exchange column throughputs.
The average concentration for each curve is obtained by integrating under
the curve. Curve 1 at 8.1 bv/hr has the lowest average effluent NHo-N
concentration (0.67 mg/l) for 150 bed volumes, but it also has the lowest
average influent NHo-N concentration (15 mg/l) . Curve 2, at 6.5 bv/hr,
has an average of 0.83 mg/1 NH^-N in 150 bed volumes of effluent with an
influent containing 17 mg/1 Nff^-N. Curve 3, at 9. 7 bv/hr, has an average
effluent NH/3-N concentration of 1.2 mg/1 and high initial NH^-N leakage,
which is due to insufficient backwash removal of residual lime remaining
in the bed after regeneration. The effluent pH for Curve 3 was 10.4- at
the time of the first sample . . Since the ammonia is poorly ionized at high
pH, the ion exchange sorption decreases and effluent NH^-N was high until
the residual lime washed out. In spite of this, NHo-N removals for 150
bed volume throughputs averaged 9^% over the three runs.
Exchange due to the 150 bed volume throughput value selected to maintain
an average NH-^-N concentration at or below 1 mg/1 uses only 55 to 58
percent of the zeolite's equilibrium capacity. The number of bed volumes
throughput per bed can be increased while maintaining low NH3~N effluent
concentrations with semi-countercurrent operation. This type of opera-
tion using two beds will be discussed later.
16
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40
en
30
20
10
0
0
D TAHOE
X POMONA
O BLUE PLAINS
I
100
200 300
MINIMUM BED VOLUMES
400
500
FIGURE 5. MINIMUM BED VOLUMES AS A FUNCTION OF NH -N CONCENTRATION
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00
o
UJ
CO
o
OPERATING CONDITIONS:
FLOW RATES:(1)8. Ibv/hr (2)6.5bv/hr, (3)9.7bv/hr
ZEOLITE GRAIN SIZE: 20x50MESH
BED VOLUME: 50 FT3
AVE. INFLUENT NH^N: (1) 15 mg/l, (2) 17 mg/l, (3) 17 mg/l
FEED: TAHOE'TERTIARY EFFLUENT
A CURVE 1
• CURVE 2
O CURVE 3
20
40
60
80
100
120
140
160
180
BED VOLUMES
FIGURE 6. AMMONIA BREAKTHROUGH CURVES FOR A 6 FT CLINOPTILOLITE BED AT VARIOUS FLOW RATES
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The effect of bed depth at a flow of 4.8 gpm/ft^ is illustrated in
Figure 7. Two 4.5 ft deep zeolite beds were run in series to provide
simultaneous data on 4.5 ft and 9 ft bed depths. The average NH^-N
conceatration in 150 bed volumes of effluent from the first bed in
series (Curve 2, 4.5 ft bed depth) was 0.81 mg/1 at a flow rate of 8.4
bv/hr. The average NHo-N concentration in the effluent from the second
zeolite bed (Curve 1, 9 ft total bed depth) was 0.35 mg/1 at a flow rate
of 4.2 bv/hr. The average NH3~N concentration in the influent was 12
mg/1.
The effect of bed depth on ammonia breakthrough with two separate
columns at 9.7 bv/hr in each case is illustrated in Figure 8. Curve
3 from Figure 6 for a 6 ft bed, and discussed previously, is repeated
and shown with data from a 3 ft bed. In general, the shallow bed of
clinoptilolite was not as effective for ammonia removal as the deep bed
at the same bed volume flow rate. The shallow bed has a lower flow
velocity which may lead to easier plugging of portions of the screen
or bed by lime or precipitated solids. Plugging would cause poor
flow distribution and lower bed efficiency.
Two Column Semi-Countercurrent Operation
Several beds in series can be operated more effectively if a column is
removed from the influent end when it becomes loaded while simultaneously
adding a regenerated column at the effluent end. This procedure moves
zeolite beds countercurrent to liquid flow. Beds can be loaded nearer
to capacity with this procedure than with single column or parallel
feed multi-column operation. The most highly loaded column is always
at the influent end backed up by one (if two in series) or more columns
having decreasing loadings and NH^-N concentrations at locations progressively
nearer the end o.f the series. Removal of a column is not decided by applying
a breakthrough criterion to the column's own effluent but by breakthrough
at the end of the series.
The performance was evaluted for countercurrent operation of three columns
(two on stream while regenerating the third). Performance data for six
runs with two columns in series are listed in Table 2 for operations at
Tahoe. The average ammonia nitrogen concentration in the effluent from
the six runs was 0.43 mg/1, and the average influent volume processed
through each column was 250 column volumes. The ion exchange columns
each contained 4.7 ft deep beds with 39 ft^ of 20 x 50 mesh clinoptilolite.
The average influent ammonia nitrogen for each run varied from 10.3 mg/1
to 16.1 mg/1, and the second column effluent varied between 0.38 mg/1 and
0.66 mg/1 ammonia nitrogen. The low and high effluent values were obtained
with the low and high influent values, respectively. Loading on the
first columns in series was terminated when the effluent from the second
columns reached 1-2 mg/1 ammonia nitrogen. Typical breakthrough curves
for the first and second columns in series with Tahoe tertiary effluent
are illustrated in Figure 9 for an average influent ammonia nitrogen
concentration of 15.1 mg/1. Ammonia loadings were increased from an
average of 57% to an average of 85% of equilibrium capacity by going from
single bed to series operation; however, the piping necessary for this is
more complicated. Ammonia nitrogen removals averaged 97%-
19
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2
OQ
OPERATING CONDITIONS
FLOW RATES: (1) 4.2 bv/hr, (2) 8.4 bv/hr
ZEOLITE GRAIN SIZE: 20x50MESH
BED DEPTH: (1)9 FT (2) 4.5 FT
BED VOLUME: (1) 76 FT^, (2) 38 FT3
AVERAGE INFLUENT NH3~N: 12 mg/l
FEED: TAHOE TERTIARY EFFLUENT
o
20
40
60
80
100
120
140
160
180
BED VOLUMES
FIGURE J. EFFECT OF BED DEPTH ON AMMONIA BREAKTHROUGH AT 4.8 gpm/ft FLOW RATE
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CD 4
UJ
CO
OPERATING CONDITIONS
ZEOLITE GRAIN SIZE: 20x50MESH
BED VOLUMES: 3 FT DEPTH = 25 FT3, 6 FT DEPTH = 50 FT3
AVE. INFLUENTNH3-N: 17
LOCATION: TAHOE
6 FT. DEPTH
20
40
60
80
100
120
140
160
180
BED VOLUMES
FIGURE 8. EFFECT OF BED DEPTH ON AMMONIA BREAKTHROUGH AT 9.7 bv/hr
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ho
to
TABLE 2
TAHOE PERFORMANCE DATA FOR SIX RUNS WITH TWO COLUMNS IN SERIES
Run No . ILoaded.
Column
1 A
2 C
3 B
4 A
5 C
6 B
Ave . Cone .
of NH3-N
in Product
mg/liter
0.325
0.42
0.38
0.43
0.66
0.44
Ave. Cone.
of NH3-N in
Influent
mg/liter
11.8
15.6
10.3
15.0
16.1
13.3
Column Vols .
Through Loaded
Column to 1.0
mg/1 Effluent
Breakthrough
290
168
311
229
215
306
Percent of
NH3-N Capacity
Used
91
62
90
83
81
102
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N3
U>
24
22
20
18
16
12
10
8
6
4
2
INFLUENT
FIRST COLUMN
EFFLUENT
SECOND COLUMN EFFLUENT
20
40
60
80 100
BED VOLUMES
120
140
160
180
200
FIGURE 9. AMMONIA BREAKTHROUGH CURVES FOR 1ST AND 2ND COLUMNS IN SERIES WITH
TAHOE TERTIARY EFFLUENT
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Performance data for 17 semi-countercurrent runs at Blue Plains with two
columns in series are listed in Table 3. The average bed volumes processed
per run was 232 and the average NHg-N in the influent and effluent was
12 mg/1 and 1.4 mg/1, respectively. An incomplete regeneration and high
feed pH in two cases caused a significant increase in the average effluent
NH3-N concentration. Excluding runs 1, 5, 12, and 14, the average NH3-N
concentration is 0.93 mg/1 for an average removal 93%. The average
bed volumes processed at Blue Plains was about 67% of the total available
capacity, which is less than that experienced at Tahoe. It is believed
that the relatively high organic concentration in the clarified raw
sewage feed stream at Blue Plains was responsible in part at least. The
differential pressure increased across the zeolite beds during operations
at Blue Plains as a result of biological growth in the beds. The growth
was effectively removed during regeneration and backwashing.
The latter runs made at each site should have been closer to steady-state.
However, there did not appear to be any definite trend in breakthrough
throughputs for specific columns at Tahoe or Blue Plains. It appears
that all semi-countercurrent loading runs were essentially at steady-state.
Zeolite Attrition
Accurate zeolite volume measurements were made at Tahoe in "A" column
only since no screen leaks developed in the column with resultant loss
of zeolite. The total zeolite volume reduction in "A" column through 15
cycles was 2.6% or 0.17% per cycle. The average ion exchange capacity
of the zeolite in all columns at the start of operations was 1.64 milli-
equivalents per gram. The average ion exchange capacity at the termination
of 15 runs per column was 1.63 milli-equivalents per gram.
24
-------
TABLE 3
PERFORMANCE DATA FOR SEVENTEEN RUNS AT
BLUE PLAINS WITH TWO COLUMNS IN SERIES
1 (a)
2
3
4
5 (b)
6
7
8
9
10
11
12 (c)
13
14 (c)
15
16
17
Ave . Flow
Rate
(gpm)
30
36
40
40
40
31
30
30
50
50
40
40
40
50
50
50
50
Ave. Cone, of
NH3-N in Influent
(mg/1)
10.0
14.0
11.0
12.3
11.6
12.5
12.2
11.5
11.4
10.6
11.8
10.3
11.7
13.6
13.3
13.4
12.7
Ave. Cone, of NH3~N
in 1st Col. Eff.
(mg/1)
2.4
4.6
6.0
6.9
3.2
6.8
5.8
1.9
4.7
5.9
4.6
4.6
5.6
3.0
5.9
5.9
6.0
Ave .
NIPj-N Cone.
of Product
(mg/1)
0.61
0.78
0.76
0.62
7.0
0.53
0.53
0.38
1.13
1.39
0.96
2.37
1.19
2.29
1.18
0.98
0.84
Bed Volumes
Treated By
Loaded Column
270
355
298
154
154
174
203
154
310
265
152
210
238
262
215
300
237
Percent of
NH3-N
Capacity Used
68
109
79
44
43
56
68
47
95
69
42
54
66
79
64
90
69
(a) two freshly regenerated beds in series
(b) second column was not completely regenerated
(c) feed pH was 10 for 1-2 hours
-------
REGENERATION STUDIES
Normal Pilot Plant Regeneration
Normal regeneration during the pilot plant studies utilized a mixture of
NaCl and CaCl2 in a solution adjusted to a pH of about 11 using lime.
The total salt concentration was generally equivalent to 0.1 N_ NaCl
which will equilibrate to a 0.02 iJ NaCl, 0.08N_ CaCl2 solution upon
continued reuse with lime addition. During regeneration the regenerant
was continuously recirculated through the zeolite bed and air stripping
column with lime addition before return to the bed. Sodium chloride
is added to the regenerant because significant amounts of sodium ion
have been found to give longer service cycles and more rapid elutions
(see Figure 5 of Appendix C). Make-up salt is added as needed to replace
that lost due to incomplete removal of regenerant following the regen-
eration cycle. High pH favors NH3 production necessary for air stripping.
A higher salt solution (total normality about 2) was subsequently chosen
to assure effective ammonia elution by a batch recycle regeneration
method using 4 bed volumes of regenerant at pH 11. Computations based
on equilibrium data indicate that ammonia elution is increased from 60%
to 88% by increasing the salt concentration of 4 bed volumes of regenerant
from 0.1 _N to 2.2 N_ at pH 11. Further increases in the salt concentration,
in particular the Ca+2 concentration, will reduce the maximum pH that
can be attained by lime addition to values significantly less than 11.
The problems due to the precipitated solids formed during regeneration
with lime solutions were quite severe at Pomona when employing normal
continuous regenerant recycle. The magnesium content of the Pomona feed
stream was approximately 20 times that of Tahoe. Examination of a zeolite
bed after regeneration at Pomona disclosed the presence of large chunks of
white material which was largely composed of magnesium hydroxide. This
material apparently filtered out on the retaining screens at the top
of the zeolite columns and periodically fell back into the bed when flow
reversals were used to reduce the pressure across the bed. A photograph
of the chunks of white material on top of a bed of clinoptilolite is
shown in Figure 10.
Both high rate backwashing with tap water and treatment with dilute
acid were temporarily effective in restoring good performance with the
zeolite beds, as illustrated in Figures 11 and 12. However, the use of
acid is expensive both with respect to the cost of the acid used and the
cost of the acid resistant piping and vessels required. High rate back-
washing was continued for each run after that for Curve 2 in Figure 11,
but was not consistent in restoring good performance. The basic problem
occurred during regeneration where a thorough elution was not always
obtained due to particulate matter blocking off portions of the bed. Removal
of the particulate matter was deemed necessary for satisfactory performance.
The effect of clarifying the regenerant before recycling through the zeolite
bed is illustrated in Figure 13. Clarification was accomplished with the
slightly inclined tube settlers in the filtration unit.
27
-------
FIGURE 10. PHOTOGRAPH OF WHITE MATERIAL ON TOP OF A
BED OF CLINOPTILOLITE AT POMONA
28
-------
INFLUENT FLOW:
INFLUENTNH3-N:
BED VOLUME:
PRIOR BACKWASH RATE:
BACKWASH:
LOCATION:
(1) AFTER LOW BACKWASH RATE
(2) AFTER HIGH BACKWASH RATE
40GPM
(1) 21 mg/J
(1) 291 GAL
(1) 20GPM
(2)
(2) 279 GAL
(2) 90GPM
TAP WATER, 2000 GAL
POMONA
20 40 60 80 100
BED VOLUMES
120
140
160
180
FIGURE 11. EFFECT OF PURE WATER BACKWASH RATE ON SUBSEQUENT NH^N BREAKTHROUGH
-------
LO
O
1 -
INFLUENT FLOW RATE: 40GPM
INFLUENT NH3-N: (1) 18 mg/£ (2) 22
BED VOLUME: 229 GAL
ACID WASH: ACETATE BUFFERED NaCI-HCI AT
LOCATION: POMONA
- (1) BEFORE ACID WASH
(2) AFTER AC ID WASH
20
40
60 80 100
BED VOLUMES
120
140
160
FIGURE 12. EFFECT OF ACID WASH ON SUBSEQUENT NH -N BREAKTHROUGH
-------
Lt
5 ~
4 -
1 -
OPERATING CONDITIONS
FLOW RATES: (1) 8.3 bv/hr (2) 8.5 bv/hr
AVE. INFLUENTNH3-N: (1) 20 mg/l (2) 19 mg/l
ZEOLITE GRAIN SIZE: 20x50MESH
BED VOLUMES: (1)39 FT3 (2) 38 FT3
LOCATION: POMONA
80 100
BED VOLUMES
FIGURE 13. AMMONIA BREAKTHROUGH CURVES FOR COLUMNS FOLLOWING REGENERATION WITH UN-
CLARIFIED (1) AND CLARIFIED (2) REGENERANT WITH HIGH MAGNESIUM CONTENT
-------
Figure 13 shows significant improvement in regenerant performance
after clarification. The combination of high rate water backwashing
with regenerant clarification should provide satisfactory regenerant
solids removal with high magnesium operations.
Pilot Plant Batch Recycle Regeneration
In order to minimize the volume of regenerant required to elute the
NH3-N from the zeolite beds, a batch regeneration technique was studied
at Tahoe. Normal regenerant was recycled through the bed only, deferring
the air stripping operation. This was done to minimize the liquid volume
to be stripped. Less liquid volume means less heat input required to
prevent freezing and keep efficiency high during cold weather stripping
operation.
Contact with several recycle batches may be necessary, however, to
obtain good regeneration. A two batch recycle regeneration could proceed
as follows. In order to start the regeneration scheme, a first batch
of fresh regenerant could be recycled through a bed followed by a second
batch of fresh regenerant recycled through the same bed. After this, the
second batch recycle used for one bed could be used for the first batch
recycle for the next bed. Each second batch recycle would consist of
fresh or renovated regenerant. Pilot plant data were collected in order
to set process specificatiors on a two batch recycle regeneration scheme.
Single batch recycle data show that the NH3-N concentration in recycled
regenerant increases rapidly from near zero to about 500 mg/1 in a few
hours with a zeolite bed loaded with an average of 2.24 g of NH3-N per
liter of bed. This loading is equivalent to ammonia removed during
reduction of a 15 mg/1 NH3-N influent to a 1 mg/1 (average) effluent
using a 160 bed volume throughput. Three elution curves are illustrated
in Figure 14 for this average bed loading using regenerant with a total
salt concentration of 0.1 N_. Curve 1 shows the lowest NH3~N (450 mg/1)
concentration after 6 hours, which is believed to be the result of the
low pH (10.8). The NH^ concentration at 450 mg/1 NH3~N and pH 10.8 is
8.9 x 10-4-M compared with 4.0 x 10~4M NH^ at pH 11.2 and 500 mg/1 NH3-N,
and 2.1 x 10-"% NH^ at pH 11.5 and 520 mg/1 NH3~N. The amount of ammonia
removed from the zeolite will, therefore, increase with pH (low solution
NH^ concentrations show low NlfJ zeolite adsorption). Curve 2 shows a
slower rate of NH3-N elution than Curve 3, which is believed due to a
combination of the slower flow rate and lower temperature and pH. Curve 3
shows that equilibrium was approached after 4.5 hours. Curves 1 and 3
represent the recycle of 2.2 bed volumes, whereas Curve 2 represents 1.2
bed volumes.
Figure 15 illustrates a first and second batch recycle for a highly
loaded zeolite bed (3.2 g/liter of bed) using regenerant with a total
salt concentration of 0.1 N. The first batch was air stripped to 107
mg/1 NH3-N while recycling through both the zeolite bed and the stripping
column. Regeneration was 75% complete at this point. The NH3-N increased
by 150 mg/1 in the regenerant after 2 hours in the second batch recycle.
32
-------
en
E
OJ
LO
700
600
500
=r 400
n:
^
2 300
3
LJ_
* 200
100
0
0
(1)
OPERATING CONDITIONS:
FLOW RATE:
PH:
BED DEPTH:
REGENERANT VOLUME:
(1)
9.4bv/hr
10.8
4. OFT
TEMPERATURE:
I i
2.2bv
14°C
j
(2)
6.4bv/hr
11.2
6. OFT
1.2bv
8°C
i
(3)
9.4bv/hr
11.5
4. OFT
2.2bv
15°C
5
TIME IN HOURS
9
10
11
12
FIGURE 14. FIRST BATCH RECYCLE ELUTION CURVES
-------
UJ
UJ
/ uu
600
500
400
300
200
100
n
FIRST BATCH ^^
RECYCLE /
A
/ OPERATING CONDITIONS:
*' FLOW RATE:
y
/ o pH:
Q^SECOND BATCH BED DEPTH:
/T RECYCLE TEMPERATURE:.
{ / BED LOADING, NH3-N:
~ / REGENERANT VOLUME:
/l l I 1 l 1 1
FIRST BATCH
6.4bv/hr
10.5-11.5
6. OFT
13 °C
3.2g/l
1.2bv
l i
SECOND BATCH
4.8bv/hr
11.4
6. OFT
19 °C
0.8g/l
1.2bv
i i
0
4
40
11
12
TIME IN HOURS
FIGURE 15. FIRST AND SECOND BATCH RECYCLE ELUTION CURVES
-------
The time required for the second batch recycle (2 hours) was less than
that of the first batch (4-6 hours) due to: l) the time lag in reaching
the optimum pH at the beginning of the first batch recycle; and 2) the
lower temperature of the first batch recycle (9°C vs. 19°C).
The regenerant system was temporarily altered at Blue Plains to permit
the recycle of 3.6 bed volumes of regenerant (iM CaCl2 and 0.2M NaCl at
pH ll). The results of this regenerant batch recycle are shown in Figure
16. The first recycle removed 75.4% of the ammonia from the zeolite bed
and the second recycle removed 12.8% with the remainder accounted for by
loss of ammonia during a flush between the two recycles. The flush con-
sisted of air stripped first cycle regenerant which was pumped through
the bed to remove regenerant that could not be drained after the first
recycle.. Essentially no regenerant could be drained from the column since
the regenerant storage vessel was on the same level with the ion exchange
vessel. The flush was air stripped to reduce the NHo-N concentration to
10 mg/1.
In addition to collecting elution data, theoretical calculations were
made. Equilibrium data were used to compute the maximum ammonia elution
from clinoptilolite as a function of regenerant volume and pH. The
results are given in Figure 17 for a pH range of 7 "to 11 using an equil-
ibrium regenerant solution at ratios of 2, 4 and 6 bed volumes of regenerant
to 1 bed volume of clinoptilolite. The initial ammonia nitrogen loading
on the clinoptilolite was 0.12 equivalents per liter or 1.7 grams per
liter. The data in Figure 17 illustrates the importance of pH in removing
sorbed ammonia during regeneration. The maximum elution values at pH 11
are higher than that shown for the first pilot plant recycle in Figure 16,
which is believed to be the result of: (l) not including potassium salt
in the recycled regenerant of the pilot plant run (potassium would build
up naturally in extended recycle service), and (2) not attaining full
equilibrium even though the pilot plant elution data indicates that a
plateau was reached.
By examining the pilot plant data, a recommended regeneration scheme has
been devised. For single column operation, a regenerant of 0.2 M NaCl,
1 M CaCl2 and adjustment to pH 11.5 with lime is recommended for two
batch recycle operation. Bed loadings will be lower in practice than used
for the pilot plant studies. It is suggested that 4 bed volume recycle
batches be used at 10 bv/hr recycle rate with 2 hours for each recycle.
This should permit a change in concentration in the first batch recycle
regenerant from 100 to 600 mg/1 NH^-N (a change of 500 was experienced
experimentally on a fresh batch) and in the second batch recycle regenerant
from 10 (from regenerant renovation) to 100 mg/1 NHo-N with almost complete
NHQ-N removal from the zeolite beds.
It should be emphasized that further practical experience may indicate
changes in regeneration procedure specifications. Calcium and sodium
concentrations were fixed rather arbitrarily in the experimental studies
35
-------
600
NH3-N LOADING: 2.4 g/£ OF ZEOLITE
FLOW RATE: 11.8 bv/hr (6.5 gpm/FT2)
400 U
CD
200 I—
FIRST RECYCLE
SECOND RECYCLE
FIGURE 16. REGENERANT BATCH RECYCLE WITH 3.6 BED VOLUMES OF
1 M CaCl2-0.2 M NaCl AT pH 11
36
-------
100
80
60
40
20
0
CONDITIONS:
.TEMPERATURE: 25°C
EQUILIBRIUM SALT CONCENTRATION
1M CaCl
0.2 M^ NaCl
0.01M KC1
REGENERANT VOLUME/
ZEOLITE VOLUME:2,4,6
8
I
9
PH
10
11
FIGURE 17. EFFECT OF REGENERANT VOLUME AND pH ON AMMONIA ELUTION
37
-------
and the regeneration procedure recommended above, although thought to
be realistic and adequate, may not be optimal. Also, the concentration
suggested for the regenerant need not be fixed at start-up. The process
could be started with 2 N_ NaCl, for example, and adjusted during operation.
This was actually done during the pilot study on electrochemical renovation
to be discussed later.
38
-------
REGENERANT RECOVERY
Air Stripping of Regenerant
The spent regenerant containing ammonia was recovered for reuse by air
stripping in a 3.6 ft. diameter by 8 ft. column packed with 1 inch
polypropylene Intalox(R) saddles. The regenerant was normally recycled
upflow through the zeolite bed at a flow rate of 4.8-7.1 gpm/ft^ until
the NH3-N approached a maximum concentration. The regenerant was then
recycled through both the zeolite bed and the air stripper until the Nt^-N
was reduced to about 10 mg/1. The liquid flow rate to the stripper was
normally 20 gpm with an air/liquid ratio of 150 cfm/gpm. Ammonia removal
in the air stripper generally averaged about 40% at 25°C. Calcium carbonate
scaling occurred on the polypropylene saddles, but did not interfere with
stripping efficiency for operating periods up to 65 days at each site. The
calcium carbonate scale formed during the Tahoe and Pomona operations was
flaky and could be removed to a large degree by spraying with water. The
scale formed during operations at Blue Plains was relatively hard and
required water fluidization of the packing to remove the scale.
The type of air stripper used in the mobile pilot plant is not recommended
for general plant use because of the energy wasted in blowing air through
the packing. A modified cooling tower with low differential pressure
across the tower is recommended.
Electrochemical Renovation of Regenerant
The chemical destruction of ammonia in regenerant solutions from the
selective ion exchange process was investigated as an alternative
method to air stripping, since atmospheric disposal of ammonia may be
undesirable in some locations. The chemical destruction of the ammonia
is accomplished by reaction with chlorine, which is generated electro-
lytically in the regenerant solution. This process can be carried out
under neutral conditions and is not as prone to solids precipitation as
alkaline lime regeneration. Regenerant solutions from the selective
ion exchange process are rich in NaCl and CaCl2 (plus MgCl2 when Mg+2 is
present in the feed stream to the clinoptilolite beds). These salts
provide the chlorine that is produced at the anode of the electrolysis
cell.
The chemical reactions that take place in the electrolysis cell are
illustrated below using NaCl as an example:
Cathode: Na+ + e~ = Na (1)
2Na + H20 = 2NaOH + H2i (2)
Anode: 2 Cl - 2e~ = C12 (3)
3C12 + 2 NH4C1= N2 + 8HC1 (4)
39
-------
The overall reaction for the destruction of ammonia with chlorine is
shown above. Excess dissolved Cl2 exists as hypochlorite (HOC1). The
hydrochloric acid produced by reaction (4) is partially neutralized
by the NaOH produced by reaction (2). However, an excess of hydrochloric
acid is produced and base must be added to the system to maintain a neutral
solution. The moles of base added should be equivalent to the moles of
NH4C1 to be reacted. The production of acid, in fact, is a good indicator
that the reactions are taking place, and the break point is clearly
indicated by stabilization of the pH.
A laboratory model electrolysis cell was obtained for conducting lab-
oratory experiments. The anode, which is the most vulnerable part of
the cell, is composed of a 5/16 inch diameter graphite rod coated with
lead dioxide. The lead dioxide coating is very resistant to attack by
chlorine or oxychloro-acids. Commercial anodes made of this material are
used in the production of chlorate and perchlorate. It is of interest
to note that ammonia impurity in the brine is destroyed early in the
production of sodium chlorate by electrolysis of brine.
Initial laboratory results with the electrolysis cell show that 5 g of
NH3-N was essentially destroyed in two hours and forty-five minutes
with the cell operating at 10 amps and 5.5 to 9 volts. A simulated
regenerant solution (1.87 N_ CaCl2, 1.31 N_ MgCl2, 0.14 N NaCl and 0.01 N_ KC1)
containing the ammonia as NH^Cl was recirculated through the cell at
flow rates of 500 to 1000 ml/min. The cell has 10 in2 of anode surface
area. At an average of 7.3 volts and 10 amps, the electrical energy
required to destroy one gram of NH3-N is 40 watt hours. When related to
the treatment of 1000 gallons of wastewater containing 15 mg/1 NH3~N
(total 57 g), the energy consumption would be 2.3 KW hours.
Although major solids precipitation was avoided, a white material was
formed on the cathode during the above experiment. Some of this material
was washed off the cathode and collected in the regenerant container.
Subsequent analysis revealed this material to be a mixture of Ca(OH)2
and Mg(OH)2- The increase in cell resistance, which required an increase
from 5.5 volts to 9 volts in order to maintain a current of 10 amps, was
due to this coating of mixed Ca(OH)2 and Mg(OH)2 on the cathode. Calcium
and magnesium are plated on the cathode where they react with water to form
their respective hydroxies. Due to the low solubility of these hydroxides,
they tend to precipitate and collect on the surface of the cathode.
Subsequent experiments with and without MgCl2 in the simulated regenerant
solutions indicate that the Mg(OH)2 increases the conductivity of the
coating but reduces its solubility in acid. The effect of Mg+2 On the
resistance of the cell is illustrated in the attached Figure 18.
Turbulent flow promoted by baffling or other cell modifications is expected
to minimize the scale formation rate. The effects of pH and flow rate on
the cell resistance were found to be insignificant. The data obtained at
pH's of 4 and 5 and flow rates of 4 and 15 1/min are shown in Figure 19.
40
-------
1 .2
1 .0
0.8
0.6
0.4
0.2
WITHOUT MgCl
REGENERANT:
TEMPERATURE
FLOW RATE:
1 .87N_ CaCl2
0.14N. NaCl
O.OIN^ KC1
40°C
4000 ml/min
-O—
WITH 0.05N_ MgCl2
10 20 30 40 50
TIME (MINUTES)
FIGURE 18. EFFECT OF Mg+2 ON CELL RESISTANCE
60
70
80
-------
-p-
1 .0
0.8
UU
0.4
o:
0.2
O pH5, 4 £/min
A pH4, 4 £/min
D pH4, 15 £/mii
REGENERANT:
1 .87N_ CaCl
0.14N, NaCl
0.05N_ MgCl
0.01N KC1
TEMPERATURE: 40°C
40
FIGURE 19.
80
200
120 160
TIME (MINUTES)
EFFECT OF H+ ACTIVITY AND FLOW RATE ON CELL RESISTANCE
240
280
-------
Preliminary pilot scale studies were conducted at the Blue Plains plant
to further evaluate electrochemical regenerant renovation. Two 500 amp
electrolysis cells and a 2000 amp, 6 volt rectifier were installed near
the mobile pilot plant for ammonia removal, along with a 20 HP pump and
a 2500 gallon regenerant treatment tank (swimming pool). Brine (NaCl)
was used as the regenerant in these pilot studies. Use of a 2 M NaCl
brine was planned but some difficulty was encountered as discussed below
and the actual concentration was 0.9 M NaCl. Regeneration of the
clinoptilolite beds in the mobile pilot plant was accomplished by pumping
brine from the regenerant storage tank (in the Recla-Mate SWB treatment
unit in the trailer) upflow through the zeolite beds to the regenerant
treatment tank. When all of the brine was collected in the regenerant
treatment tank, recirculation of the brine through the electrolysis cells
was started. Based on laboratory data, the two 500 amp cells are capable
of destroying about 150 grams of ammonia nitrogen per hour in chemically
pure solutions. Chlorine is produced at the surface of the lead dioxide
anode in the cells where it reacts with the solution to produce hypochlorite
Chlorine is produced at the rate of 657 g per hour when the total current
applied to the cells is 1000 amps.
Since both the Recla-Mate SWB unit and the regenerant treatment tank each
holds only half of the required volume of regenerant, it was necessary
to put the regenerant through each bed twice. Data on elution with neutral
regenerant is included in Appendix B. Ammonia destruction was accomplished
after each transfer except for the second batch of the final run. Ammonia
elution from the clinoptilolite beds did not appear to be as rapid as that
expected for a 2 M solution. Although sufficient salt (2000 Ibs NaCl) was
added to the SWB unit to make a 2 H NaCl solution, subsequent analysis
of the regenerant showed that it contained only 0.9 moles per liter. It
is believed that some concentrated brine was lost through the filter of
the SWB unit during the dissolution step. Since the conductivity of
the solution in the electrolysis cells was satisfactory, no significant
loss was suspected at the time.
Three zeolite beds were regenerated with the brine; however, only two
of the beds were loaded with ammonia. The first bed was assumed to be
partially loaded after regeneration with a lime-salt solution, but was
subsequently found to contain very little ammonia as evidenced by the low
NH3-N concentrations in the brine during the first run.
The highest NH3~N levels were attained in the first batches of brine from
the second and third runs and were 248 mg/1 and 210 mg/1 for 1800 gal and
2100 gal, respectively. The difference in regenerant volumes was due to
incomplete draining of the column during the second run. Gassing in the
zeolite bed caused by release of nitrogen from the destruction of ammonia
was suspected as the cause of the incomplete draining, since a significant
amount of hypochlorite was present in the brine after the first run.
Gassing was also observed in the laboratory when a clinoptilolite bed
was eluted with brine containing a relatively high concentration of hypo-
chlorite. The electrical energy consumed by the first batches from the
second and third runs was 5.67 x 10? coulombs and 4.60 x 10? coulombs,
43
-------
respectively. The amount of energy per gram of NH3~N was 33,600 coulombs
and 27,500 coulombs with a power consumption of 54 and 46 watt hours/g,
respectively. The energy and power consumed in the second batch of brine
in the second run was 22,000 coulombs and 35 watt hours per gram, respectively.
The second batch contained 72 mg/1 NE^-N. The electrolysis cells were
operated at 1000 amps and 5.8 to 6.0 volts for the most part. No significant
increase in cell resistance was noted, but the differential pressure across
the cells increased with time as a result of scale formation. The
difference in electrical energy consumption may be due to a difference
in the amount of organic matter in the regenerant, although this has
not been determined yet. The zeolite beds were given a short backwash
to remove slime from the bed prior to regeneration and some variation in
the amount of this material remaining after backwash may have existed
between the two runs. Clarified raw sewage was used to load the zeolite
beds at Blue Plains and this may have caused the slime. The slime
problem will not be significant if the zeolite columns are operated
after carbon columns which sorb the organic nutrients causing the slime.
Breakpoint chlorination for the first batch of regenerant in Run 3
is illustrated in Figure 20. The total NH^-N remaining in the brine
after 13 hours of treatment was 0.16 mg/1. The nitrate-nitrogen present
in the brine was 8.0 mg/1. The latter represents the total produced by
the two runs or a 1.7% conversion of NH3~N to N03~N. No ammonia was
detectable by direct nesslerization at the breakpoint of each batch.
The Ca"1"^ increased from 1900 mg/1 in 1800 gal after the second regeneration
to 2700 mg/1 in 2100 gal after the third regeneration. Lime was used
to neutralize the acid formed by the destruction of NH^Cl.
Ammonia leakage was relatively high after regeneration with the 0.9 M
NaCl. The high leakage (1-3 mg/1 NH3-N) is attributed largely to high
pH in the zeolite beds during the initial part of the service cycle. The
high pH was apparently caused by residual lime left in the system by
previous lime regenerations. The service cycle times were 37 hours
and 34 hours (318 bed volumes and 292 bed volumes, respectively) on the
two loaded columns regenerated with electrolytically renovated brine.
44
-------
100
80
C\J
o
ct:
o
60 —
40 -
20 -
1000 AMPS
6.0 VOLTS
30-35°C
6-
2100 GAL
210 mg/1
TOTAL CURRENT:
VOLTAGE:
TEMPERATURE:
pH:
REGENERANT VOLUME
INITIAL NH3-N:
13
TIME, HOURS
FIGURE 20. BREAKPOINT CHLORINATION OF REGENERANT BATCH #1, RUN 3
-------
ACKNOWLEDGMENTS
The studies reported herein were conducted by personnel of the Pacific
Northwest Laboratories, otherwise known as Battelle-Northwest, a division
of Battelle Memorial Institute. Basil W. Mercer and Ronald G. Arnett
served as Technical Director and Pilot Plant Engineer, respectively.
The efforts and assistance of the following Battelle personnel contri-
buted to the successful completion of the project. Pilot plant operations:
Neil Bonney, James A. Coates, Marvin J. Mason, Richard G. Parkhurst,
R. Gregory Swank and Robert G. Upchurch. Bench scale and theoretical
studies: John G. Adams, David A. Cochran, Gaynor W. Dawson, and R. Jeff Serne,
Technical and administrative assistance: Lloyd L. Ames, Gordon L. Gulp,
Douglas E. Olesen, Alan J, Shuckrow and C. Joseph Touhill.
Our sincere appreciation is extended to the following individuals outside
Battelle whose cooperation and assistance made this work possible.
Mr. Russell L. Gulp of the South Tahoe Public Utility District, South Lake
Tahoe, California; Mr. David R. Evans of Cornell, Rowland, Hayes, and
Merryfield, Engineers and Planners, Corvallis, Oregon; Messrs. Jerry C.
Wilson and Harlan E. Moyer of Glair A. Hill and Associates, Consulting
Engineers, Redding, California; Mr. Charles W. Carry of the Sanitation
Districts of Los Angeles County, Los Angeles, California; Mr. John N.
English of the Environmental Protection Agency, Water Quality Office,
Pomona, California; Mr. Gerald Stern of the Environmental Protection
Agency, Water Quality Office, Los Angeles, California; Mr. Dolloff F.
Bishop of the Environmental Protection Agency, Water Quality Office,
Washington, D. C.; and Mr. Alan Cassel of the District of Columbia
Sanitation District, Washington, D. C.
The support of the project by the Environmental Protection Agency, Water
Quality Office and the assistance and suggestions provided by Dr. Robert
B. Dean, WQO Project Officer, are gratefully acknowledged. Also appreciated
are the technical and editorial suggestions made by Dr. Harry Bostian of
the Environmental Protection Agency during review of the report.
-------
REFERENCES
1. Slechta, A. F. and G. L. Gulp "Water Reclamation Studies at the
South Tahoe Public Utility District". Journal Water Pollution
Control Federation, p. 787 (1967).
2. Johnson, W. K. and G. J. Schroepfer. "Nitrogen Removal by Nitrifi-
cation and Denitrification". Journal Water Pollution Control Federa-
tion, p. 1015 (196^)-
3- Ames, L. L., Jr. "Zeolitic Removal of Ammonium Ions from Agricultural
and Other Wastewaters". Proceedings of the Thirteenth Pacific North-
west Industrial Waste Conference, Washington State University, Pullman,
Washington (April, 1967).
ik Mercer, B. W.; Ames, L. L.; Touhill, C. J.; Van Slyke, W. J.; and
Dean, R. B. "Ammonia Removal from Secondary Effluents by Selective
Ion Exchange". Journal Water Pollution Control Federation, ^2, Part 2,
R95, (1970).
5. Glasstone, S. "Textbook of Physical Chemistry," Second Edition, New
York. D. Van Nostrand Company, Inc. 19^6, p. 956-59.
-------
APPENDIX A
SAMPLE CALCULATION OF AMMONIUM ION
LOADING USING ACTIVITY COEFFICIENTS
The experimentally measured selectivity coefficient is defined as
K
A
(B) (A)a
s z
(A)a (B)b
s z
(A-l)
where
(A),
(B) = normality of cations A and B in equilibrium solution
S
(A)z> (B)z = equivalent fractions of cations A and B on the zeolite
a, b = the number of cations A and B represented in the chemical
reaction for exchange of A and B.
A "thermodynamic"selectivity coefficient incorporating activity coefficients
can also be defined:
'A
thermo
b x_x b
Y (B)
(A-2)
The y's are activity coefficients that correct for solution non-ideality
and the "thermodynamic" selectivity coefficient is therefore a constant
value for solutions having different concentrations of ions but the same
ratio of A to B. It is assumed that corrections are not necessary for
ions adsorbed on the zeolite. The invariance in the thermodynamic
coefficient serves as the basis for a procedure to correct experimental
values and calculate equilibrium zeolite loadings.
The experimental selectivity coefficients were determined in a reference
solution of 0.1 _N in cations. The thermodynamic selectivity coefficient
can be expressed in terms of this reference solution:
A
thermo
b /T,Nb , , a
B (B)S (A)Z
a /, \ a , -.b
A (A)S (B)z
0.1H Cad,
/ b
,1 N CaCl,
K;
0.1 N. CaCl,
(A-3)
50
-------
Since the thermodynamic selectivity coefficient is independent of
reference solution, the above specific expression can be equated to
the general expression giving,
(I
A
0.1 N CaCl,
0.1 N CaCl,
B <»>*
Rearranging, the B ion loading can be expressed in terms of the A ion
loading
(B)
(YB/Y*) do*
-D JTX O
b, a\
B A 0.1 N CaCl
Z
(A)!
.£.
a RB
s A o.i
N CaCl0
— 2
(A-4)
Activity coefficients may be neglected in univalent-univalent exchange,
but are necessary in univalent-divalent exchange. The activity coefficients
can be predicted using Debye-Huckel theory. A sample calculation using
the above expression to predict equilibrium zeolite loading is shown next.
To calculate the ammonium ion loading on clinoptilolite, the fact that
the equivalent fractions of all the ions on the zeolite must sum to
one is used:
(Ca)
(Mg)
(A-5)
where (A) represents the equivalent fraction of the cation exchange
capacity of the zeolite occupied by ion A.
From typical simulated secondary effluent (SSE) specifications in Table A-l
= 4.75
(K)
= 0.317
= 2083
(NVs
= 1458
where (A) = normality of ion A in solution.
s
From Debye-Huckel theory , the activity coefficient of ion i is:
,2,
.(5)
= -0.509
I = 1/2 ZM± Z±2 , I is ionic strength
51
-------
TABLE A-l
SIMULATED SECONDARY EFFLUENT, SSE
Cations Concentration, mg/1 Normality
Na+ 130 5.7 x 10~3
K+ 15 3.8 x 10~4
NH* 20 1.2 x io~3
Ca+ 60 3.0 x 10~3
Mg"1"1" 25 2.1 x 10~3
52
-------
M^ = concentration gram ions/liter of ion i
Z. = charge on ion i
for SSE
Na+ = 5.70 x 10~3
K+ = .38 x 10~3M
NH4' = 1.20 x 10~3M
4
i i o
Ca = 1.50 x 10 M (1/2 normality)
i i Q
Mg = 1.05 x 10~ M (1/2 normality)
- -3
Cl = 12.4 x 10 M (assuming that Cl is the only negative ion)
for SSE
I = 0.0149
Using the Debye-Huckel relationship
log1Q YCa++ = -0.509 (2)2 A/0.0149
Yp -H- = 0.564
(-13.
Iog10 ^H4. = -0.509 (I)2
^H4" = 0.868
4
For ~0.1 N_ CaCl2 solution used to determine selectivity coefficients
(NH, in the equilibrium solution can be neglected),
I = 1/2 [(0.05) (2)2 + (0.1) (I)2]
I = 0.15
2
Iog1() rCa = -0.509 (2)'
Yo ++ = 0.162
Iog10 YNHt = -0.509 (I)2 A/0.
YNH~I" = 0.635
4
53
-------
SSE
= 0.747
0.41
1.83
0.1 N CaCl,
This factor of 1.83 is the correction factor between the SSE solution
and the 0.1 N solution used to determine the values of the selectivity
coefficient,
Since Mg has the same charge as Ca , we will assume that the correction
factor for K|&» is also 1.83. Univalent-univalent exchanges will not
require an activity correction. Using equation A-4,
Na
0.1 N CaCl,
0.1 N CaCl,
= 11, (Na)z = 4.75
= 0.33, (K) = 0.317 (NH,)
Z 0.33 4
0.1 N CaCl,
= 760 , (Ca) = (2083) (1.83) (NH )'
Z 760 4
0.1 N CaCl,
= 2400, (Mg) = 1458 (1.83) (NH.)
Z 2400 4
Substituting in Equation A-5 ,
(NH4)z +5.02
+1.11
+0.96
+0.43 (NH ) = 1 ,
2.39 (NH. ) +6.13 (NH. ) -1=0,
4 z 4 z
(NH )
4 z
-2.39 + A/5.71 + 2452
= 0.254
12.26
Since 1.81 meq/g is the total capacity of the zeolite
(.254) (1.81) = 0.46 meq/g
compared to 0.4 meq/g determined experimentally.
54
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APPENDIX B
PRELIMINARY DESIGN OF A 10 MGD AMMONIA REMOVAL
PLANT UTILIZING ELECTROLYTIC RENOVATION OF SPENT REGENERANT
design of a plant using electrolysis for removal of ammonia from
spent regenerant must be considered preliminary at this time due to
ack of pilot scale data needed to optimize the operating parameters.
service flow rates are the same as that (6 bv/hr) used in South
toe's design of an air stripping lime regeneration process in Appendix C;
/ever, this flow rate is believed to be very conservative. No problems
poor flow distribution because of solids plugging associated with the
2 of lime are anticipated. This is because the electrolysis process
.1 be operated at neutral conditions. Laboratory data has indicated
at flow rates up to 20 bv/hr can be used without significant leakage or
emature breakthrough.
e capital cost of this installation is minimized through the use of
enforced concrete tanks for the ion exchange beds rather than closed
:eel tanks as specified in South Tahoe's design. Further, open concrete
inks will be acceptable with respect to ammonia volatilization in the
Lectrolysis process since very little ammonia will volatilize from the
sutral regenerant solution. Where lime solids are a problem, the steel
ressure tanks noted in Appendix C offer an advantage of quick removal
nd cleaning of distribution screens. Fouling of the gravel underlayer
f an open concrete tank would pose a serious problem of cleaning.
he ion exchange units are patterned after those used for water softening
.t the LaVerne Filtration Plant of the Metropolitan Water District of
iouthern California. The LaVerne units have been operated successfully
)ver a period of thirty years for softening Colorado River water. The
Largest ion exchange units presently used in the LaVerne Plant have a
zross section of 1590 square feet. The ion exchange beds in this design
nave a cross section of 800 square feet. This ammonia removal process
employs four zeolite beds each containing 3200 cu ft of 20 x 50 mesh
clinoptilolite. Three beds will be in service at any given time while
the fourth bed will be undergoing regeneration. The zeolite beds will
be operated in parallel to 150 bed volumes each before removal from service
for regeneration. The service flow direction will be downflow and regen-
eration will be accomplished upflow. The relatively high density of the
wet clinoptilolite granules (1.7 g/cm^) will prevent localized fluidization
or channeling during the upflow regeneration. The design specifies 15
volumes of neutral regenerant, which contains two equivalents per liter of
mixed calcium, sodium, magnesium and potassium chlorides, for eluting the
ammonia from the beds at 23°C. Laboratory elution data for a similar
solution are given in Figure B-l and shows almost 90% elution after 15
volumes of regenerant throughput. The evolution of heat during the
55
-------
electrolysis step will increase the temperature of the regenerant, which
is expected to increase the ammonium ion elution rate. However, no high
temperature elution data are available to confirm a significant reduction
in regenerant volume or regeneration time.
The regeneration step will require about 7.5 hours at 2 bv/hr. Regenerant
will be pumped from the regenerant storage tank through the bed to the
regenerant processing tank. At the end of the regeneration cycle, a
rinse step will be initiated to push the remaining regenerant out of the
zeolite voids to the regenerant processing tank. The regenerant in
the processing tank will be pumped through a bank of electrolysis cells
which generate chlorine for destruction of the ammonia. A flow diagram
of the regeneration processing step is given in Figure B-2. A total of
530 cells, operating at about 6 volts and 500 amps, will be employed in
the electrolysis of the spent regenerant. A lime slurry will be fed to
the cell effluents to maintain a neutral pH. The regenerant will be pumped
through an enclosed tank (50,000 gal) to capture the hydrogen evolved
in the process. The hydrogen is vented to the atmosphere or burned.
Approximately 7 cu ft of hydrogen gas will be generated from the destruction
of ammonia in 1000 gallons of wastewater at an NH3-N concentration of
15 mg/1. A 10 MGD plant will therefore be producing about 70,000 cu ft
of hydrogen per day. The cost of recovering the hydrogen, including
scrubbing to remove the nitrogen trichloride, is more than the thermal
value of the hydrogen.
Figure B-3 illustrates the design of the reinforced concrete tanks for
the zeolite beds. The beds are contained in tanks measuring 40 ft in
length, 20 feet in width, and 8.25 ft in depth. The zeolite beds are
4 ft deep and are supported on 12 in of 4 layers of graded gravel. The
total height of 8.25 ft allows for backwash and the distrbution system.
The distribution system over the beds for regenerant draw-off and
service feed consists of a trough running down the middle of the tank
lengthwise with connecting troughs with overflow weirs running laterally,
spaced 4 feet apart. The distribution system under the beds consists of
the perforated concrete units of the type employed at the LaVerne plant.
Summaries of costs for a 10 MGD plant for treating tertiary effluent
and using electrochemical renovation of the regenerant are shown in
Tables B-l and B-2. The estimates are patterned after those used by
South Tahoe in Appendix C except where electrochemical renovation or the
modifications discussed above require changes. One exception is that
a flat 46% of capital is used to estimate auxiliaries, contingenices
and engineering design. This is roughly equivalent to similar expenses
which are itemized in Appendix C. Less make-up clinoptilolite is
required than in the design in Appendix C because neutral operation
will eliminate solids precipitation and the need for backwashing with
the accompanying zeolite loss. Electrical energy for electrolysis was
assumed to be 50 watt hours/g NH3-N. The costs for lime and NaCl are
the same as for the South Tahoe design. Although the chemistry is different,
the net amount of lime required for neutralization is the same. Make-up
NaCl is required to compensate for losses; there is no net consumption by the
chemical reactions in either process.
56
-------
1 .0
o
X
8 10
COLUMN
12 14
VOLUMES
18 20
22
FIGURE B-l.
ELUTION OF HECTOR CLINOPTILOLITE, 30 g,
20-50 MESH, LOADED WITH 1. 2N_ NH/jCL + 0.83 N_
KC1, XQ = 37.3 MEQ NHj/30 g COLUMN. SEVEN
COLUMN VOLUMES/HR ELUTION RATE. ELUTING
SOLUTION CONTAINED 0.1425 N NaCl + 0.0095 N_
KC1 + 1.8747 N_ CaCl2 + 1.3122 N^ MgCl2-
X/X0 = FRACTION ELUTED.
57
-------
H0 DISPOSAL
Ln
OO
REGENERANT
PROCESSING
TANK
REGENERANT
STORAGE
TANK
ELECTROLYSIS
CELLS
RECTIFIER
FIGURE B-2. REGENERANT PROCESSING FLOWSHEET
-------
FEED
INLET'
ON-OFF VALVE
FLOW CONTROL VALVE
ON-OFF VALVE
SPENT REGENERANT
GRAVEL
ON-OFF VALVE
REGENERANT BRINE INLET
-C&J-
REGENERANT DISTRIBUTOR
HEAD
ON-OFF VALVE
PRODUCT WATER OUTLET
FLOW CONTROL VALVE
FIGURE B-3. TYPICAL VALVE ARRANGEMENT FOR ION EXCHANGE TANK
-------
TABLE B-l
PRELIMINARY CAPITAL COSTS - ELECTROLYSIS PROCESS - 10 MGD
Description
Electrolysis unit complete with cells,
rectifier, piping and bus bars with
20% extra cells
Ion exchange beds with distribution
system
Regenerant processing and storage
tanks — 500,000 gallon
Hydrogen recovery tank
Piping
Instrumentation
Valves
Regenerant processing pumps—
15,000 gpm
Product water pumps—4,000 gpm
Regenerant-backwash pumps—6,000 gpm
Zeolite
Plus 46% auxiliary, contingencies and
engineering design
Quantity
636 cells
Total Cost
$ 420,000
48,000
96,000
1
-
-
30
3
3
3
12,400 cu ft
Equipment Total
28,000
139,000
79,000
77,000
17,000
7,000
9,000
124,000
$1,044,000
480,000
Total Capital Cost$l,524,000
60
-------
TABLE B-2
PRELIMINARY ESTIMATE OF TOTAL COST FOR ELECTROLYSIS PROCESS
Cost/MG
Lime $ 6.50
Make-up Sodium Chloride 6.90
Make-up Clinoptilolite 0.20
Chlorine 6.40
Electricity 42.80
Anode Replacement (2 yr life) 4.40
Operational Labor 14.00
Maintenance, Material and Labor 9.20
Total Operating Cost 90.40
Capital Amortization, 6% for 20 yr 36.40
Total Cost $126.80
61
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APPENDIX C
ENGINEERING DESIGN OF A 7.5 MGD AMMONIA REMOVAL PLANT
UTILIZING AIR STRIPPING FOR RECOVERY OF SPENT REGENERANT
This appendix presents an engineering design report prepared by the
South Tahoe Public Utility District, South Lake Tahoe, California.
Experimental data used in the design, and discussed in preceding sections,
were obtained when Battelle's mobile pilot plant was located at South
Lake Tahoe. This design was for a plant utilizing air stripping to
remove ammonia from the clinoptilolite regenerant. The report is
appended here to combine designs with alternate approaches to regenerant
recovery in a single publication.
The Environmental Protection Agency (EPA) should be substituted for the
Federal Water Quality Administration (FWQA) in the following report.
62
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PHASE I ENGINEERING DESIGN REPORT
SUPPLEMENTING AMMONIA STRIPPING WITH FURTHER
NITROGEN REMOVAL BY SELECTIVE ION EXCHANGE
AND BREAKPOINT CHLORINATION
by
Jem- C. Wilson and David R. Evans
Clair A. Hill & Associates Cornell, Howland, Hayes & Merryfield
Consulting Engineers Engineers & Planners
Redding, California 96001 Corvallis, Oregon 97330
Project Director: Technical Consultant:
Russell L. Gulp, Manager Harlan E. Moyer
South Tahoe Public Utility District Clair A. Hill & Associates
South Lake Tahoe, California 95705 Consulting Engineers
Redding, California 96001
for the
FEDERAL WATER QUALITY ADMINISTRATION
Program #17010EEZ
Contract #14-12-561
FWQA Project Officer, Dr. R. B. Dean
Advanced Waste Treatment Research Laboratory
Cincinnati, Ohio
April, 1970
Project No. L-145.98
63-
-------
ABSTRACT
Pilot plant investigations of the efficiency of nitrogen removal from a tertiary
treated sewage effluent were conducted at the South Tahoe Public Utility District.
The nitrogen removal process utilized ion exchange and employed a natural zeolite,
clinoptilolite, which is selective of ammonium ions in the presence of sodium,
magnesium, and calcium ions. Regeneration of the exhausted clinoptilolite is
accomplished with solutions or slurries containing lime. Lime provides hydroxyl ions
which react with the ammonia ions to yield an alkaline aqueous ammonia solution.
The ammonia solution is processed through a heated air stripping tower to remove
the ammonia which is exhausted harmlessly to the atmosphere. The regenerant
solution is not discarded and the process generates no liquid wastes. Ammonia
removals as high as 99% can be obtained using this process.
The pilot plant investigation provided the design criteria on which a preliminary
design of a IVz mgd plant is based. The plant design provides a system utilizing 12
exchange beds, nine of which are in service at all times, and three of which are in a
regenerant cycle. Each bed would be 12 feet in diameter and have an effective
clinoptilolite depth of 8 feet. The flow through the bed would be 6 bed volumes per
hour or 680 gpm.
The construction and operating costs of the 7/2 mgd plant are estimated. Based on
current costs in the Lake Tahoe region, the process is estimated to cost $84.95 per
million gallons to operate, and capital costs amortized at 6% interest for 20 years are
estimated to be $63.10 per million gallons.
This report was submitted in fulfillment of Contract 14-12-561 (17010 EEZ)
between the Federal Water Quality Administration and the South Tahoe Public
Utility District.
Key Words: Nitrogen Removal, Ion Exchange, Clinoptilolite, Ammonia Stripping,
Tertiary Treatment Costs.
65
-------
TABLE OF CONTENTS
SECTION PAGE
INTRODUCTION 71
PROPOSED PROCESS AND DESIGN CRITERIA
Process Description 75
Breakpoint Chlorination 78
Design Criteria 79
PILOT PLANT STUDIES
Service Cycle 81
Elution Cycle 84
Ammonia Stripping Cycle 88
Temperature Requirements for Elution & Air Stripping 91
THE PROPOSED PLANT
Service Cycle 93
Ion Exchange Beds 93
Clinoptilolite Transfer Tank 93
Regenerant Cycle 94
Elutrient Storage Tanks 94
Lime Storage & Feeders 95
Sodium Chloride Storage & Brine Feeder 95
Lime and Salt Mixing Basin 95
Ammonia Stripping Tower & Recycle Basin 96
Process Controls 96
Heating of Elutrient and Stripping Tower Air 98
Sludge Collection 98
FINANCIAL REQUIREMENTS
Construction Cost Estimates 101
Incidental Cost Estimates 101
Operating Cost Estimates 101
Total Project Cost 102
REQUESTED GRANT PROGRAM
General 105
Proposed Schedule 1°6
Project Grant Costs 107
67
-------
LIST OF TABLES AND FIGURES
TABLE OR
FIGURE NO TITLE PAGE
TABLE A TYPICAL CARBON COLUMN EFFLUENT
QUALITY BEFORE CHLORINATION 76
TABLE B PROPOSED DESIGN CRITERIA 80
FIGURE 1 AMMONIA BREAKTHROUGH CURVES
FOR A 6 ft CLINOPTILOLITE BED
AT VARIOUS FLOW RATES 82
FIGURE 2 AMMONIA BREAKTHROUGH CURVES
FOR TWO 4.5 ft CLINOPTILOLITE
BEDS IN SERIES 83
FIGURES FIRST BATCH RECYCLE ELUTION CURVES 85
FIGURE 4 FIRST AND SECOND BATCH RECYCLE
ELUTION CURVE 86
FIGURE 5 EFFECT OF SALT ADDITION ON NH3
ELUTION DURING BATCH RECYCLE 87
FIGURE 6 PERCENT AMMONIA REMOVAL VS CUBIC FEET
OF AIR PER GALLON WASTEWATER TREATED
FOR VARIOUS DEPTHS OF PACKING 8 9
FIGURE 7 PERCENT AMMONIA REMOVAL VS SURFACE
LOADING RATE FOR VARIOUS DEPTHS
OF PACKING 90
TABLE C ESTIMATED CAPITAL COSTS 103
TABLE D ESTIMATED OPERATION COSTS 104
68
-------
LIST OF DRAWINGS
DWG NO (Drawings are bound at end of the report)
0 VICINITY MAP & LOCATIONS PLAN
1 TOTAL PLANT FLOW DIAGRAM
2 PROCESS FLOW DIAGRAM
3 MECHANICAL PLANS
4 SECTIONS & DETAILS
5 SECTIONS & DETAILS
69
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INTRODUCTION
The potential benefits of removing nitrogen from wastewaters are well known and
will not be recited here in detail. The principal benefits from nitrogen removal to a
low residual include restriction of algal growths, and elimination of the toxic effects
of ammonia nitrogen on fish and aquatic life. Ammonia removal also enhances the
efficacy of chlorination in disinfection of wastewaters.
Today, advanced wastewater treatment processes are developed to the extent that
practical, reliable, and plant-proven methods are available to produce reclaimed
water of such high quality that it may be used for any desired purpose.
The costs are reasonable too, with two notable exceptions, the costs for nitrogen
removal and for dewatering of sewage-chemical sludge mixtures. The latter problem
can be avoided by separate settling and handling of sewage sludge and chemical
sludge in those situations where the mixture is too expensive to dewater. Also, it is
possible that methods may be developed to cut the cost for handling difficult sludge
mixtures.
Unfortunately, there does not appear to be in the offing a solution to the high cost
of nitrogen removal. Ammonia stripping is by far the cheapest means of nitrogen
removal which has, at this time, a high degree of reliability and ease of control.
However, ammonia stripping is subject to considerable loss of efficiency at water
and air temperatures near freezing, and it is not practical to operate the process at all
at air temperatures below freezing.
Deposition of scale on stripping tower packing is a problem at present, but it appears
probable that this problem can be solved. Assuming solution of the calcium
carbonate deposition problem, which appears likely, then ammonia stripping may be
a good method for nitrogen removal in warm climates. It could be supplemented by
breakpoint chlorination where required to remove residual ammonia following
stripping.
In cold climates it is undoubtedly still economical to use ammonia stripping at air
temperatures above freezing, provided it can be supplemented at temperatures
slightly above freezing and supplanted entirely by another method of nitrogen
71
-------
removal at air temperatures below freezing. This project was undertaken with this
approach in mind. The existing ammonia stripping tower was to be supplemented or
supplanted, according to air temperatures, by the selective ion exchange process for
the removal of ammonia nitrogen from wastewater as developed to the pilot plant
stage by Battelle-Northwest under contract to FWQA. The process employs a natural
zeolite, clinoptilolite, which is selective of ammonium ions in the presence of
sodium, magnesium, and calcium ions. Regeneration of the exhausted clinoptilolite
is accomplished with solutions or slurries containing lime. Lime provides hydroxyl
ions which react with the ammonium ions to yield an alkaline aqueous ammonia
solution. This ammonia solution is processed through a heated air stripping tower to
remove the ammonia which is exhausted harmlessly to the atmosphere. The spent
regenerant is then fortified with more lime and recycled to the zeolite bed to remove
more ammonia. Since the regenerant is not discarded, the process generates little
liquid wastes. Ammonia removals as high as 99% can be obtained using this process.
The last 0.5 to 1.0 mg/1 of ammonia nitrogen can then be removed by use of
breakpoint chlorination. Battelle-Northwest has constructed a 100,000 gpd mobile
demonstration plant for further study of this process. They very kindly conducted
tests over a period of several months last year at the Tahoe plant. The results of their
pilot plant work on the Tahoe reclaimed water established the design criteria for the
full-scale plant on which this report is based.
Phase I of the contract with FWQA consists of the preparation of an engineering
design report. Phase II is the construction of full-scale plant facilities using selective
ion exchange for supplemental nitrogen removal. Phase III covers collection,
analysis, and reporting of data on the efficiency and cost of nitrogen removal by ion
exchange, ammonia stripping, and breakpoint chlorination, for a 2-year operational
period.
This report completes Phase I of the project, and gives a basis for consideration of
proceeding with Phase II. It presents the design criteria developed by Battelle-
Northwest in their pilot plant work as applied to a preliminary design of a full-scale
plant for ammonia removal by selective ion exchange. The description of the
proposed plant is illustrated by drawings showing the general layout and design of
the facilities. Cost estimates are also presented. The proposed plant has been
designed for a nominal capacity of 7H mgd, which corresponds to that of the
existing advanced wastewater treatment plant. The decision to build a T^-mgd plant
addition is based upon several important considerations. First, it has been observed
72
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that most of the great number of visitors who tour the plant and export facilities, as
well as local residents, judge the accomplishments of the plant almost solely on the
quality of the final effluent as represented by the appearance of the Indian Creek
Reservoir. If only part of the plant flow is treated, then the quality of the reclaimed
water as produced by the plant does not truly reflect the process capabilities. This is
not just an idea, it is a fact which has already been driven home as a result of
experience with the ammonia stripping tower which has a capacity of only one-half
that of the rest of the plant. This means even during low flow seasons, that part of
the plant flow must be bypassed during maximum hourly flows. Test data can be
gathered under these conditions, but the value of plant scale demonstration is
completely lost, or even becomes negative, so far as visitor impressions are
concerned. We are of the firm opinion that a successful plant scale demonstration
absolutely requires that the capacity of all facilities be built to full 7Vi-mgd capacity.
There is more than visitor impressions involved. Indian Creek Reservoir is being
intensively studied so far as algal growths are concerned. Since the reservoir is
drained down to about one-third full capacity each summer for irrigation purposes,
it is possible to test under natural conditions the effects of nitrogen removal.
Actually, phosphorus reductions to a residual of about 0.1 mg/1 have given excellent
control of algal growths in the reservoir. Against this background of field data,
further possible benefits from nitrogen removal can be evaluated.
73
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PROPOSED PROCESS AND DESIGN CRITERIA
Process Description
The selective ion exchange process developed by Battelle-Northwest11' for removal
of ammonia nitrogen from wastewater is the basis for a proposed full-scale plant at
the South Tahoe Water Reclamation Plant.
The process employs a natural zeolite, clinoptilolite, which is selective for
ammonium ions in the presence of sodium, magnesium, and calcium ions.'2' A high
pH lime solution containing NaCl and CaCl2 is used to regenerate the exhausted
clinoptilolite. The solution provides both ions for exchange with the ammonium
ions and hydroxyl ions to yield an alkaline aqueous ammonia solution. The
following equation approximates this reaction:
2NH4+1R-! + Ca(OH)2 — Ca+2R2-! + 2NH3 + 2H2O
The presence of a significant amount of sodium on the clinoptilolite lengthens the
service cycle and shortens the elution cycle.
After the ammonium ion is eluted from the clinoptilolite, the high pH alkaline
aqueous ammonia solution is passed through an air stripping tower where the
ammonia is stripped from the regenerate or elutrient. Make-up lime and salt are
added to replace the exchanged calcium and sodium.
At the South Tahoe Water Reclamation Plant, the proposed clinoptilolite exchange
process would be added to the existing plant following the carbon adsorption step as
shown on Drawing No. 1. The quality of the water, see Table A, is such as to
preclude organic fouling of the ion exchange beds.
A schematic diagram of the proposed ion exchange beds, lime elutrient system, and
ammonia air stripping system is shown on Drawing No. 2. For design flows, nine
beds would be in service and three beds in regeneration.
The direction of flow for the beds in service would be downflow. All beds would
operate in parallel.
(1) Developed by Battelle Memorial Institute, Pacific Northwest Laboratories Division, under contract with
the Federal Water Pollution Control Administration.
(2) Battelle-Northwest, "Research Report, Ammonia Removal from Agricultural Runoff and Secondary
Effluents by Selective Ion Exchange to the FWPCA," December 1968, 56 p.
75
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Table A
Typical Carbon Column Effluent Quality
Before Chlorination
BOD (mg/1) 1 2
COD(mg/l) 5 15
Turbidity (JU) 0.3 - 0.8
MBAS(mg/l) 0.1-0.2
pH 7.0-8.0
Coliforms (mpn/ 100ml) less than 50
Nitrogen (mg/l-N)
Ammonia 4-20
Nitrite 0.1 - 0.6
Nitrate 1 - 2
Suspended Solids (mg/1) -0-
Total Dissolved Solids (mg/1) 250 - 350
Phosphorus (mg/1 PO4-P) 0.05 - 0.15
Alkalinity (mg/1 CaCO3) 150 - 250
Hardness (mg/1 CaCO3) 50-175
Chlorides (mg/1) 25
Sulfates (mg/1) 25
Carbon Fines some present
76
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When a given volume of carbon column effluent has passed through a set of three
beds, for example, beds 1, 2, and 3, this set of beds would be taken off line for
regeneration. At this time, from a previous regeneration, elutrient tank A would
contain elutrient water with a very high ammonia content (say 600 mg/1); tank B
would contain elutrient water with a low ammonia content (say 100 mg/1); and tank
C would contain ammonia-free elutrient water (say 10 mg/1). The contents of tank A
would be air stripped during the regeneration of exchange beds 1, 2, and 3. The
regeneration would proceed as follows:
1. Exchange beds 1,2, and 3 would be drained to the final effluent.
2. Low ammonia content elutrient water from tank B (100 mg/1) would be
recirculated upflow through the three exchange beds and back through tank B to the
exchange beds until the concentration of ammonia in the elutrient began to
approach a maximum value (say 600 mg/1). Throughout the recirculation, make-up
rime and salt would be added. A pH of about 11.5 would be maintained.
3. After an allotted time (long enough for elutrient from tank B to approach
a maximum ammonia concentration, the elutrient would be changed to recirculation
to and from tank C through beds 1, 2, and 3. Tank C with its ammonia-free elutrient
water would be recirculated for an allotted time (long enough for elutrient from
tank C to reach about 100 mg/1). At this stage of the elution, the small amount of
ammonia left on the zeolite would be distributed uniformly throughout the bed.
Tank A with ammonia-free water (10 mg/1 NH3, water stripped during the
regeneration of beds 1, 2, 3) would be pumped once upflow through the bed to
further polish the lower portion of the bed and prevent leakage of ammonia during
the downflow service cycle.
The elutrient in tank B (600 mg/1 NH3) would be held for air stripping during the
regeneration of the next set (say beds 4, 5, and 6) of ion exchange beds. Tank C
with 100 mg/1 elution water would become the lead tank for this next set of ion
exchange beds. Tank A with ammonia-free elution water (10 mg/1 NH3, water
stripped during the regeneration of beds 1, 2, 3) would be used as the polishing tank
for beds 4, 5, and 6.
4. Once the elution of beds 1, 2, and 3 was completed, the three beds would
be drained back to tank A.
77
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5. Beds 1, 2, and 3 would then be filled slowly (from the bottom to remove
trapped air) with product water from the other nine beds in service.
6. After the beds were filled with product water, more product water would
be pumped at a high rate through beds 1, 2, and 3 in sequence. The backwash water
would be routed through the sludge line to the floe basin, which is located just
ahead of the chemical clarifier in the existing plant. (See Drawing No. 3)
7. After backwashing was completed, ion exchange beds 1, 2, and 3 would
be placed in service and beds 4, 5, and 6 would be taken offline for regeneration.
Ammonia in the elutrient solution would be removed by air stripping at a pH of
about 11.5. In the preceding example, during the regeneration of beds 1, 2, and 3,
the very high ammonia content (600 mg/1) in the elutrient solution of tank A was to
be air stripped. The following procedure would be used:
1. The contents of tank A would pass through the tower down into the
recycle basin below the tower.
2. The contents of the recycle basin would then be pumped back up through
the tower once again. This time, however, the effluent from the tower would flow
back to tank A.
3. The contents of tank A would now contain about 10 mg/1 of ammonia,
and would be ready to serve as the polishing volume during the regeneration of ion
exchange beds 4, 5, and 6;
Breakpoint Chlorination
In order to remove the last 0.5 1.0 mg/1 NH3-N remaining after treatment with
clinoptilolite, breakpoint chlorination would be accomplished just prior to pumping
the reclaimed water to Luther Pass. No new equipment is needed for this purpose, as
the existing chlorinators are adequate.
78
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Design Criteria
Phase I of the District's research contract with the FWQA includes the collection of
pilot plant data using the Battelle-Northwest mobile pilot plant and carbon column
effluent from the reclamation plant. The data are used in preparing this report.
A summary of the design criteria based on the pilot work accomplished September
through December 1969 is listed in Table B.
79
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TABLE B
PROPOSED DESIGN CRITERIA - AMMONIA REMOVAL BY
SELECTIVE ION EXCHANGE AND BREAKPOINT CHLORINATION
AT THE SOUTH TAHOE WATER RECLAMATION PLANT
Capacity 7.5 mgd
Ammonia Concentration
Influent 15-20 mg/1 NH3-N
Ion Exchange Effluent 0.5-1.0 mg/1 NH3-N
Chlorinated Effluent 0 mg/1 NH3-N
Exchange Beds
Length of Service Cycle 150 bv W
Service Cycle Loading 6 bv/hour
Bed Depth 8 ft
Bed Diameter 12 ft
Exchange Bed Regeneration
Quantity of Elutrient 8 bv
Elution Rate 10 bv/hr
Elutrient (Initial - Ca build-up during service)
Calcium Oxide 500 mg/1
Sodium Chloride 2.0 N
pH H.5
Ammonia Stripping Tower
Capacity 300 gpm
Air/Water Ratio 300 cfm/gpm
Hydraulic Loading 3.5 gpm/ft2
Packing Height 24 ft
Efficiency per Pass 85%
Number of Passes 2
Temperature
Elutrient 740
Stripping Tower Air 740
Breakpoint Chlorination 10 mg Cl2/mg NH3-N
(1) bv means gross bed volumes
80
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PI LOT PLANT STUDIES
Service Cycle
Pilot scale studies on ammonia removal from South Tahoe Public Utility District
tertiary effluent by selective ion exchange were conducted over a 4-month period
with a 100,000-gpd mobile pilot plant. This plant has three 500-gallon ion exchange
vessels (3.25 ft in diameter by 8 ft high) which were filled with 38 ft3 (284 gal) to
50 ft3 (374 gal) of the natural zeolite, clinoptilolite. Feed water was percolated
through the zeolite beds either singly or in series until ammonia breakthrough
occurred. The beds were regenerated with solutions containing lime, NaCl and
CaCH. The spent regenerant containing ammonia was recovered for reuse by air
stripping the ammonia from the regenerant solution in a 3.6 ft diameter by 8 ft
column packed with one-inch polypropylene Intalox (R) saddles. The flow rate to
the stripping column was normally 20 gpm with an air/liquid ratio of 150 cfm/gpm.
Ammonia removal in the air stripper averaged about 40% at 25°C. The spent
regenerant was normally recycled through the air stripper and zeolite bed until the
NH3-N concentration in the stripped regenerant was reduced to 10 mg/1.
Zeolite ammonia loading studies were carried out to establish the volume of feed
water that can be processed through a zeolite bed until significant ammonia
breakthrough occurred. The computed NE^-N loading on two 4.5 ft beds of zeolite
(9 ft total bed depth) with South Tahoe Public Utility District tertiary effluent
containing 12 mg/1 NH3-N was 6.52 g of NH3-N per gallon of zeolite. The average
NH3-N concentration in the effluent was approximately 0.4 mg/1 at a flow rate of
4.2 bed volumes per hour. On this basis the volume processed for the 9-ft bed was
150 bed volumes. The volume selected for full-scale plant design is 150 bed volumes
at a flow rate of 6 bed volumes per hour with a 12-foot-diameter bed, 8 feet in
depth. An average effluent NH3-N concentration of 0.5 to 1.0 mg/1 is expected
under these conditions, with 15-20 mg/1 NH3 in the influent.
Figure 1 illustrates NH^-N breakthrough curves for a single 6-foot bed of zeolite
operated at flow rates varying from 6.5 to 9.7 bed volumes per hour with 15 to 17
mg/1 NH3-N in the influent. Curve 1 at 8.1 bed volumes per hour shows the lowest
average effluent NH3-N concentration (0.67 mg/1) to 150 bed volumes but also has
the lowest average influent NH3-N concentration (15 mg/1). Curve 2 at 6.5 bed
volumes per hour shows an average of 0.83 mg/1 NH3-N in 150 bed volumes of
81
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00
to
OPERATING CONDITIONS:
FLOW RATES:(1) 8.1 bv/hr (2) 6.5 bv/hr, 6) 9.7 bv/hr
ZEOLITE GRAIN SIZE: 20x50MESH
BED VOLUME: 50 FT3
AVE. INFLUENTNH3-N: (1) 15 mg/l, (2) 17 mg/l, (3) 17 mg/l
FEED: TAHOE TERTIARY EFFLUENT
FIGURE 1
A CURVE 1 AMMONIA BREAKTHROUGH CURVES
FOR A 6ft. ClINOPTILOUTE BED
A PIIRVF ?
9 ^ur\vL t AT VAR|OUS FLOW RATES
O CURVE 3
20
40
60
80
100
120
140
160
180
BED VOLUMES
-------
00
Q
UJ
CO
OPERATING CONDITIONS
FLOW RATES: (l)4.2bv/hr, (2)8.4bv/hr
ZEOLITE GRAIN SIZE: 20x50 MESH
BED DEPTH: (1)9 FT (2) 4.5 FT
BED VOLUME: (1) 76 FT^, (2) 38 FT3
AVERAGE INFLUENT NH3-N: 12 mg/l
FEED: TAHOE TERTIARY EFFLUENT
o
FIGURE 2
AMMONIA BREAKTHROUGH CURVES
FOR TWO 4.5ft. CLINOPTILOLITE
BIDS IN SERIES
(2)
20
4U
60
80
100
120
140
160
180
BED VOLUMES
-------
effluent with an influent containing an average of 17 mg/1 NH3-N. Curve 3 at 9.7
bed volumes per hour shows an average effluent NH3-N concentration of 1.2 mg/1
NH3-N. Curve 3 also shows an initial high NH3-N concentration which is the result
of insufficient backwash removal of residual lime remaining in the bed after
regeneration.
Ammonia-nitrogen breakthrough curves for two 4.5-foot zeolite beds operated in
series are illustrated in Figure 2. The average NH3-N concentration in 150 bed
volumes from the first bed in series (Curve 2) was 0.81 mg/1 at a flow rate of 8.4 bed
volumes per hour. The average NE^-N concentration in the effuent from the second
zeolite bed (Curve 1) (total 9-foot bed depth) was 0.35 mg/1 at a flow rate of 4.2
bed volumes per hour. The average NH3-N concentration in the influent was 12
mg/1.
It is evident from the data above and from series operation data that the higher
influent NH3-N concentrations will produce higher effluent NH3-N values.
Breakthrough and leakage are influenced by bed packing characteristics and are
subject to variation, other factors being equal. In general, deep beds at low flow
rates will yield the low leakage and sharp breakthrough curves.
Elution Cycle
In order to minimize the volume of regenerant required to elute the NH3-N from the
zeolite beds, a batch regeneration technique was studied. A minimum volume of
regenerant is desired to maintain a low heat input during the air stripping operation
in cool weather. Pilot plant data show that the NH3-N concentration in recycled
regenerant increases rapidly from near zero to about 500 mg/1 in a few hours with a
zeolite bed loaded with an average of 8.5 g of NH3-N per gallon of bed. Three
elution curves are illustrated in Figure 3 for average bed loading of 8.5 g of NH3-N
per gallon of bed. Curve 1 shows the lowest NH3-N (450 mg/1) concentration after 6
hours, which is believed to be the result of the low pH (10.8). The NH+4
concentration at 450 mg/1 NH3-N and pH 10.8 is 8.9 x 10"% compared with 4.0 x
10-4M NH+4 at pH 11.2 and 500 mg/1 NH3-N and 2.1 x 10-4M NH+4 at pH 11.5
and 520 mg/1 NH3-N. The amount of ammonia removed from the zeolite will,
therefore, increase with pH (low solution NH+4 concentrations show low NH+4
zeolite adsorption). Curve 2 shows a slower rate of elution of the regenerant solution
than curve 3, which is believed due to a combination of the slower flow rate and
lower temperature and pH. Curve 3 shows that equilibrium was attained after 4.5
hours.
84
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00
FIGURE 3
FIRST BATCH RECYCLE
ELUTION CURVES
OPERATING CONDITIONS:
FLOW RATE:
PH:
BED DEPTH:
REGENERANT VOLUME:
(1)
9.4bv/hr
10.8
4. OFT
2.2bv
TEMPERATURE:
i i
I
14°C
(2)
6.4bv/hr
11.2
6. OFT
1.2bv
8°C
(3)
9.4bv/hr
11.5
4. OFT
1.2bv
15°C
10
11
12
TIME IN HOURS
-------
00
UJ
/ uu
600
500
400
300
200
100
n
FIRST BATCH -x^
RECYCLE/'* FIRST
FIGURE 4
AND SECOND
BATCH
>^ RECYCLE ELUTION CURVE
^
/ OPERATING CONDITIONS
•/ FLOW RATE:
/ o P*
Q^SECOND BATCH BED DEPTH:
/T RECYCLE TEMPERATURE:
{ / BED LOADING, NH3-N:
/ REGENERANT VOLUME:
/ l l l 1 l 1 1
: FIRST BATCH
6.4bv/hr
10.5-11.5
6. OFT
13 °C
3.2g/l
1.2bv
l i
SECOND BATCH
4.8bv/hr
11.4
6. OFT
19 °C
0. 8 g/l
1.2bv
i i
0
8
10
11
12
TIME IN HOURS
-------
CO
260
240
220
200
180
160
14O
120
100
80
^ fin
E 60
Z
m
I"
-n20
O
c
•a
m
01
•X*
t
i
/
y
/,
v
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r
Salt
Added
•^ r
EF
/>
^s'
:FECT
>
/
/
OF SAL
Dl
/
V
•)
f
.TADD
JRING
P
/
OPE RATH
Cl
IGURE
ITION(
BATCH
TE
pH
RE
BE
SA
5
DN NH;
RECY(
SJG cor
OW RA
MPERA
GENEF
D VOL
LT AD[
|-N ELL
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4DITIO
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ANT V
JME
DED
ITION
*JS:
OLUME
i
22°C
11.2
2 3 bv
240^1
10lbs
91 23436789 10
TIME IN HOURS
-------
Curves 1 and 3 represent the recycle of 2.2 bed volumes; whereas, curve 2 represents
1.2 bed volumes. In order to attain a sufficiently high NH3-N removal, the recycle of
4 bed volumes of regenerant is recommended at pH 11.5 with regenerant containing
0.2M Na+ plus 1M Ca+2 which is in equilibrium with the Na+. At 500 mg/1 removal
in the 4 bed volumes of recycled regenerant (average loading of 10 g of NH3-N per
gallon of zeolite), the first batch recycle will elute 75% of the loaded NH3-N.
The second batch recycle is expected to bring the elution up to more than 90% and
will contain about 100 mg/1 NH3-N. Since the second batch recycle is used without
air stripping for the first batch recycle of the next regeneration, it then will attain an
ultimate NH3-N concentration of 600 mg/1 at 75% elution. Figure 4 illustrates a first
and second batch recycle for a highly loaded zeolite bed (3.2 g of NH3-N per liter of
bed). The first batch was air stripped to 107 mg/1 NH3-N while recycling through
both the zeolite bed and the stripping column. Regeneration was 75% complete at
this point. The NH3-N increased by 150 mg/1 in the regenerant after 2 hours in the
second batch recycle, which indicates that the regeneration will be close to 100%
with 4 bed volumes. The time required for the second batch recycle (2 hours) was
less than that of the first batch (4-6 hours) due to : (1) The time lag in reaching the
optimum pH at the beginning of the first batch recycle and; (2) the lower
temperature of the first batch recycle (9°C vs 19°C).
The effect of salt (NaCl) addition during a batch recycle is illustrated in Figure 5.
The salt was added when the NH3-N concentration in the regenerant appeared to be
leveling off at 155 mg/1. The NE^-N concentration then increased at a rate higher
than that prior to the salt addition. The loss in regenerant per cycle is about 5% of a
bed volume based on laboratory data which is approximately 1.3% of 4 bed volumes
of regenerant. The total salt lost per thousand gallons of water treated is estimated
to be 0.32 Ibs of NaCl from a regenerant solution containing mixed salts (CaCl2 +
NaCl) at a concentration of 2.2 equivalents per liter. The maximum pH that can be
attained by adding lime to this regenerant is 11.5.
Ammonia Stripping Tower
The design criteria for the ammonia stripping tower were based on pilot plant
studies by the South Tahoe Public Utility District/1) Figures 6 and 7 indicate that
85% removal at 74°F can be achieved with an air-to-liquid ratio of 300 cfm/gpm and
a surface loading of 3.5 gpm/ft2 in a 24 ft tower.
(I) Smith and Chapman, "Recovery of Coagulant, Nitrogen Removal, and Carbon Regeneration in Waste
Water Reclamation." FWPCA Grant WPD-85, June 1967.
88
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100
24' DEPTH
<
O
LJ
O
IT
LLJ
Q.
80
^ 12 FEET OF PACKING
20 FEET OF PACKING
h 24 FEET OF PACKING
200
400
600
800
1000
1200
CUBIC FEET AIR/GALLON TREATED
FIGURE6
PERCENT AMMONIA REMOVAL VS CUBIC FEET
OF AIR PER GALLON WASTEWATER TREATED
FOR VARIOUS DEPTHS OF PACKING
89
FIGURES
-------
<
>
o
tr
<
H
O
cc
1 2 FEET OF PACKING
20 FEET OF PACKING
24 FEET OF PACKING
1.0
2.0 3.0 4.0 5.0
SURFACE LOADING RATE - gpm/ft2
FIGURE?
PERCENT AMMONIA REMOVAL VS SURFACE LOADING RATE
FOR VARIOUS DEPTHS OF PACKING
6.0
7.0
c
33
m
-------
An open packing such as found in cooling towers was chosen over other types of
packing such as Intalox^ saddles or Raschig rings in order to avoid high, expensive
pressure losses and to utilize existing pilot plant data.
Temperature Requirements for Elution and Air Stripping
The choice of 23°C (74°F) as the elution water and air stripping temperatures was
based on laboratory and pilot plant work by Battelle-Northwest. It was felt there
was not enough data to accurately predict elution and air stripping requirements for
lower or higher temperatures.
L.L. Ames showed that at 23°C (74°F) the elution rate probably represented an
optimum between Ca(OH>2 solubility in the eluting solution and the cation
exchange kinetics.'1) He reported that zero degrees and 80°C represented the
slowest elution rates, requiring 80 to 120 bed volumes.
If the laboratory work by Ames at low temperatures were used to design the
proposed South Tahoe plant, the elution rate would be at least twice as long. For
the proposed design flows at the longer elution rate, 15 ion exchange beds instead of
12 would be needed. The size of each exchange bed would be about the same. Nine
beds would be in service and six in regeneration. Each elution tank and the tower
recycle basin would have to be about 100% larger than required for 12 exchange
beds.
The flow rate to the stripping tower would be about the same, however, for a
24-foot-high tower, about four times JT ore surface area would be needed at the
cooler temperatures. The fan would have to deliver nearly four times as much air.
The larger tower would also require greater maintenance expense for removal of
calcium carbonate.
(1) Ames, L.L. "Zeolitic Removal of Ammonium Ions From Agricultural Wastewaters," 13th Pacific
Northwest Industrial Waste Conference Proceedings, 135,152, Washington State University, April 1967.
91
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-------
THE PROPOSED PLANT
Service Cycle
At a design capacity of 7.5 mgd, nine ion exchange beds would be in service at all
times. The beds would be operated in parallel. When a total of 450 bed volumes
(3.05 mg) had passed through three beds (150 bv per bed), these beds would be
taken off line and three freshly regenerated beds would be put on line in their place.
A flow totalizer would be used to measure the 150 bed volumes and automatically
initiate an elution cycle.
Ion Exchange Beds
Drawing No. 4 shows a typical cross-section of an ion exchange bed. Each bed would
be 12 feet in diameter and have an effective clinoptilolite depth of 8 feet. The flow
through a bed would be six bed volumes per hour or 680 gpm. The surface loading
rate would be 6 gallons per square foot per minute. Overdesign is about 20% due to
choosing an integer (ft) diameter.
Due to calcium carbonate deposition during regeneration with pH 11+ lime solution,
the inlet and outlet screens for each bed are removable for cleaning. The screens can
be taken out of the bed without removing the clinoptilolite.
Tne beds have also been designed such that clinoptilolite can be added or removed
by means of a water/clinoptilolite slurry. A special transfer header is included to aid
in moving the clinoptilolite out of the bed and into the transfer line.
Clinoptilolite Transfer Tank
The transfer tank would be used for both washing and adding make-up clinoptilolite.
To add make-up clinoptilolite, about 5% of an exchange bed volume could be
dumped from bags into the transfer tank, washed if necessary, and then transferred
to the top of any of the exchange beds in a slurry by pressurizing the transfer tank
with water.
If clinoptilolite in an exchange bed required washing, it would be removed from the
bottom by pressurizing an exchange vessel, put into slurry form, and moved to the
transfer tank. After washing, the clinoptilolite would be returned in a slurry to the
top of the exchange bed.
93
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Regeneration Cycle
At design flow the service cycle for a set of three beds would be about 25 hours. On
the basis of nine beds in service and 25-hour service cycle, a set of beds would have
to be regenerated about every 8 hours.
i
Under the elution procedure outlined earlier, the elution of ammonium ions would
take place in two phases. Four bed volumes (81,600 gal.) would pass through a set
of three beds for each phase; or a total of 8 bed volumes would be needed to
completely elute the ammonium ion from each of the three exchange beds. Each
elution phase would last about 4 hours. The elution would be upflow at a rate of 10
bed volumes per hour or 10 gallons per square foot per minute.
Two of the three 1,700-gpm pumps (Drawing No. 3) would be needed to provide
enough flow to elute three exchange beds in parallel. The three pumps would be
rotated so that one could be offline in order to remove lime build-up. After elution,
one of the three pumps would be used to backwash the bed. Each bed would be
washed upflow in sequence at a rate of 15 gpm/ft^.
It is anticipated that most of the pipes, valves, and pumps, particularly those items
which handle the lime elutrient, would have to be routinely cleaned. To facilitate
cleaning of the pipes, all changes in direction would be made by using crosses with
blind flanges. Wafer stock valves would be used in order that the pipe cleaning device
would readily pass through the valve.
The regeneration cycle, as was the case for the service cycle, will be completely
automatic.
Elutrient Storage Tanks
Three covered concrete elutrient storage tanks, each holding 94,500 gallons, or
approximately 20% more than the anticipated elutrient volume, would be used.
The elutrient tanks would also function as settling basins for excess calcium
carbonate. The tank floors have a l-to-12 slope to a sludge draw-off sump. A slotted
pipe is used to remove the elutrient in order to avoid disturbing the settled sludge. In
the same manner, elutrient returning from the exchange beds flows into a stilling
well before entering the tank.
94
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Approximately a 4-hour period would elapse from the time ammonia-free elutrient
returned from the stipping tower to the elutrient tank until it was needed for an
elution polishing cycle. The return flow from the stripping tower would enter at the
back of the tank and be used to move the settled solids toward the sludge sump. The
4-hour delay would allow time to clarify the elutrient. The rectangular tank design
would permit, if necessary, the addition of mechanical sludge collection equipment.
Lime Storage and Feeders
At a design flow of 7.5 mgd, theoretically 2,300 Ibs/day of calcium oxide would be
required for elutrient make-up if the ion exchange beds were to remove 19 mg/1
NH3-N. If a 40% safety factor is included and the lime the District can purchase has
a CaO content of 90% approximately 3,600 Ibs of lime per day would be needed.
The proposed storage bin was sized to provide 10 to 14 days of storage.
About one-half of the ammonia is eluted from the exchange beds in the first hour
regeneration. To meet this high demand for calcium hydroxide in the first hour, the
lime feeder and slaker would have to be capable of handling about 800 Ibs/hr. Two
800 Ibs/hr feeder-slakers would be needed to allow for down-time for cleaning and
maintenance.
Sodium Chloride Storage and Brine Feeder
Sodium chloride, 0.2 N, was found to be helpful during the elution of ammonia
from the clinoptilolite. To make up 0.2N NaCl in the elutrient, 4,700 Ibs/day would
be needed, or on the basis of the proposed 2-week storage period, 66,000 Ibs of
sodium chloride would have to be stored.
Lime and Salt Mixing Basin
The mixing basin would be used for both the addition of calcium hydroxide and
sodium chloride makeup and for initial clarification of the elutrient. The elutrient
velocity through the basin would be about 0.1 ft/sec and the detention time about 9
minutes. The baffles within the basin would provide for 60 feet of linear flow
distance. A hopper bottom would permit sludge draw-off.
95
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Automatic pH monitoring equipment at the inlet and outlet to the basin would be
used to control the lime feeder. The brine feeder would be paced by flow in the
basin.
Ammonia Stripping Tower and Recycle Basin
The ammonia stripping tower has been sized to treat the contents of an elutrient
tank in 8 hours, using two passes through the tower at 85% removal per pass. At a
flow rate of 300 gpm, an air-to-water ratio of 300 cfm/gpm, and a loading of 3.5
gpm/ft2, approximately 86 ft2 of packing and 90,000 cfm air would be required/1)
To prevent loss in stripping efficiency due to calcium carbonate build-up, the tower
packing would be removable in sections for cleaning. The catch basin below the
packing is sloped to the recycle basin below the tower to aid in sludge removal.
The design of the recycle basin below the tower would be similar to that of the
elutrient tanks. A sloping bottom and sludge draw-off are included. The design is
such that mechanical sludge collection could be added, if necessary, at a later date.
Process Controls
The selective ion exchange process has three basic operations; service cycle,
regeneration cycle, and air stripping of the high ammonia regenerant. These
operations occur simultaneously during normal operation of the process. The zeolite
beds have four headers, bed influent and bed product, regenerant in and out, each
header requiring a valve at each bed. During normal operation, nine beds are in
service and three beds are in regeneration; when three beds are regenerated, they
must be placed on line and three more beds taken off line for regeneration. This
requires the opening or closing of 24 valves, plus the regenerant storage tank valves.
While the bed regeneration is proceeding, the regenerant from the previous
regeneration must be passed through the ammonia stripping tower twice. It is
obvious the three basic operations must be automated with pneumatic or
hydraulically actuated valves.
" ' On the basis of processing 4 bv regenerant/requiied ion exchange bed volume, the tower is overdesigned
by about 9%. On the basis of actual bed volume about 11% underdesign is indicated but this is an unrealistic
basis because the beds are overdesigned. The tower should be more than adequate as specified.
96
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In addition, to maintain the three-bed-regeneration sequence, a flow totalizer must
be placed on a minimum of every third bed, and preferably on every bed, to record
the number of bed volumes passed before regeneration is necessary. Also, each bed
should have a flow rate recorder to maintain equal flows to each of the three beds.
After the three beds are regenerated, product water from the nine beds in service
must be diverted and pumped through the three beds for backwashing. A flow
controller must be placed on the product water to throttle the flow and divert the
required supply of water for backwashing. Also, valves should be placed on the
present carbon column effluent line or bed influent header and on the product to
chlorination line to provide for bypassing of the ion exchange process, if necessary.
The regeneration backwash pumps and the tower pumps must be automated in one
control center, since both elution and stripping operations will be proceeding at the
same time.
Information obtained during pilot plant operation indicated the optimum
temperature and pH to achieve maximum efficiency of the elution and stripping
operations were at least 74°F and a pH of 11-11.5. Also, the presence of a
significant amount of sodium increases the service cycle and shortens the elution
cycle. In order to maintain optimum conditions, a pH monitoring system must be
provided on the entrance and exit of the mixing basin to control lime slaker
operation. The salt brine addition equipment must be automated in accordance with
the flow through the mixing basin, to maintain an approximate .2 molar
concentration in the regenerant. Temperature monitoring and controls must be
placed immediately up stream of the zeolite beds and the stripping tower to
maintain the optimum temperature in the regenerant for elution and stripping. In
addition, the air utilized by the stripping tower must be maintained at the optimum
temperature of 74°F. The air temperature monitoring and heat addition equipment
must be automated to maintain optimum temperature conditions.
The accumulation of solids in the mixing basin, regenerant storage tanks, and the
tower recycle basin will require daily pumping of sludge to the existing lime floe
basin. Sludge pumping and the necessary valving can be accomplished manually.
In the pilot plant studies there was evidence of the zeolite "mudballing" with the
cohesive solids in the regenerant. Manually operated valves have been provided to
remove the zeolite from the bed, transfer the zeolite in a slurry to the transfer tank,
wash it, and transfer it back to the bed.
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During the initial start-up of the process, a large number of ammonia-nitrogen
determinations will be required to determine a more accurate time period for the
two elution cycles. When approximate time periods are established, the regeneration
cycle can be placed on an automatic time basis, and should eliminate laboratory
analysis except for periodic checks. Also, ammonia-nitrogen determinations will be
required initially to verify the expected efficiency of the ammonia stripping tower,
and the 150 bed volume service cycle before one mg/1 ammonia-nitrogen
breakthrough. After these design criteria are verified only the routine plant
operation analysis will be required, except for periodic checking of decreased
efficiency of the stripping tower as a result of calcium carbonate incrustation on the
tower fill. Periodically the fill will be removed and cleaned.
Heating of Elutrient and Stripping Tower Air
In order to maintain both the elutrient and stripping tower air at 74°F, five separate
heating systems would be needed.
The design has been based on utilizing the 500-gpm scrubber water off the exising
lime and sludge furnaces to preheat the air to the tower. The remainder of the air
heating would be accomplished with 180°F hot water from a new boiler. The heat
would be transferred to the air by means of coils occupying a 7-foot-wide by
11-foot-high area on each end of the building opposite the open tower packing.
Three water-to-water heat exchangers would be used to heat the elutrient. Heat
exchanger No. 1 would reheat to 74°F the elutrient going to the top of the stripping
tower from either the elutrient storage tank or from the stripping tower recycle
basin. Heat exchanger No. 2 would reheat the elutrient after it had made the second
pass through the stripping tower but before it reached the elutrient storage tanks.
The third heat exchanger, No. 3, would heat the elutrient to 74°F as it passed from
a storage tank to an ion exchange bed during the elution cycle.
Sludge Collection
Ammonia removal studies were conducted with the mobile demonstration plant at
the Richland wastewater treatment plant. From these studies it has been estimated
that 0.4 Ibs of dry solids per 1,000 gal of treated water would be generated. For the
proposed process at the South Tahoe Water Reclamation Plant, approximately 3,300
Ibs of dry solids would be produced per day.
98
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The sludge would be pumped from the chemical mixing basin, each elutrient tank,
and the stripping tower recycle basin through a 12-inch pipe to the existing lime
flocculation basin. The 12-inch pipe would also be used to carry 1,700 gal per
minute of backwash water from the final ion exchange bed rinse.
One 530-gpm sludge pump, operating for 15 minutes per day, would be used to
pump the sludge to the floe basin. The high pump rate would be necessary in order
to maintain scouring velocities in the 12-inch pipe. A second 530-gpm pump would
be provided as a stand-by unit.
A new lime transfer pump at the existing chemical clarifier and a new centrifuge
would be needed to handle the additional sludge.
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FINANCIAL REQUIREMENTS
This section presents the estimated project costs and a brief summary of the
required project financing.
Construction Cost Estimates
The construction cost estimates are based on the quantities and character of work
described for each element of the project. Materials prices were obtained from
manufacturers and were added to estimated labor costs plus contractors' overhead
and profit, to obtain individual costs. These prices were then compared to actual bid
prices for similar work, adjusted as necessary to reflect past experience in
construction contract costs for the Lake Tahoe area. The estimates are based on
construction starting in 1970. If the project is delayed then the cost estimates must
be escalated to account for probable increases in material and labor costs. Table C is
a detailed estimate of the project construction cost.
Incidental Cost Estimates
An allowance for incidental costs must be added to the construction cost. These
costs are an integral part of the project cost and include engineering, construction
inspection, administration of the project, and collection and publishing of operating
data after completion of the project.
Operating Cost Estimates
Operation of the project will lead to increased costs for the District. Table D shows
the estimated operating cost for the nitrogen removal plant when it is in complete
operation. These costs are based on a design flow of 7.5 mgd and will be somewhat
higher per millon gallons for flows less than the design capacity, since some of the
costs are fixed. Labor costs are based on one operator per shift, 3 shifts per day, 365
days per year. Chemical costs are based on predicted dosages required for
continuous operations. All prices are based on delivery to Tahoe. Electric power and
natural gas costs are based on current rate schedules at Tahoe.
101
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Total Project Cost
The total estimated construction cost, including engineering and inspection, is
$1,981,500. The average annual capital cost, amortized over 20 years at 6% interest,
would be $172,700 per year. Based on a plant capacity of 7.5 mgd the capital cost
per million gallons would be $63.10. The estimated annual operating cost is
$232,100.
The data collection and publication phase upon completion and operation of the
project is estimated to cost $72,000 per year, based on two engineers being assigned
full time to the project.
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TABLE C
ESTIMATED CAPITAL COSTS
DESCRIPTION
QUANTITY UNIT UNIT COST TOTAL COST
Move-in
Bonds £ Insurance
Excavation & Backfill
Concrete
Steel Building
Painting
Welded Steel Pipe
Large & Special Valves
Misc. Fabricated Steel
Floe Basin
Ion Exchange Beds
Clinoptilotite
Lime Storage & Slakers
Salt Storage & Feeder
Salt Conveyor
Booster Pump
Regenerant Pump
Sludge Pump
NH3 Tower Pump
NH3 Tower Structure
Boiler & Accessories
Air Heat Exchanger
Water Heat Exchangers
Hot Water Pumps
Electrical Work
Instrumentation
Heating & Ventilating
Misc. Piping & Plumbing
Operational Tests
Cleanup
1
12
12,000
2
3
2
2
LS
LS
LS
LS
LS
LS
LS
LS
LS
ea
ea
cf
LS
LS
LS
ea
ea
ea
ea
LS
LS
LS
LS
LS
LS
LS
LS
LS
LS
LS
$30,000
27,000
10
4,500
5,000
1,500
1,500
TOTAL ESTIMATED CONTRACT PRICE
Construction Contingencies
Design Engineering at 6.37%
Construction Inspection ($3,000/mo x 18 mos)
TOTAL ESTIMATED CAPITAL COST
$ 5,000.00
2,000.00
15,000.00
165,000.00
65,000.00
25,000.00
275,000.00
60,000.00
25,000.00
30,000.00
324,000.00
120,000.00
30,000.00
30,000.00
20,000.00
9,000.00
15,000.00
3,000.00
3,000.00
50,000.00
24,000.00
8,000.00
8,000.00
4,000.00
100,000.00
200,000.00
15,000.00
10,000.00
5,000.00
2,000.00
$1,647,000.00
165,000.00
$1,812,000.00
115,500.00
54,000.00
$1,981,500.00
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TABLE D
ESTIMATED OPERATION COSTS
AT 7.5 mgd
COST/YR COST/mg
Makeup Lime $ 19,900 $ 7.30
Makeup Sodium Chloride 18,900 6.90
Makeup Clinoptilolite 53,600 19.60
Operational Labor 50,800 18.60
Maintenance, Material & Labor 25,000 9.15
Chlorine 17,400 6.40
Natural Gas 20,300 7.40
Electricity 26,200 9.60
Total Operating Costs $232,100 $ 84.95
Amortized Capital Cost 172,700 63.10
(6% - 20 yr.)
TOTAL CAPITAL & OPERATING COST $ 148.05
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REQUESTED GRANT PROGRAM
General
The removal of nitrogen from sewage effluent remains as one of the most difficult
problems to solve in the advanced waste treatment field. The various methods that
have been tried at South Tahoe, both as pilot plant studies and full-scale operation,
have not accomplished the complete removal of nitrogen that is necessary if sewage
disposal in the Lake Tahoe basin is ever to be considered. The need for a complete
nitrogen removal process is, of course, a national problem, not limited to Lake
Tahoe only.
The proposed ion exchange process as envisioned in this report, gives promise of
being a most satisfactory means of complete nitrogen removal. The pilot plant work
that has been completed strongly indicates that ion exchange is effective, but also
indicates that the process will be relatively expensive and also will be relatively
sophisticated to properly operate.
The only assured means to adequately demonstrate the effectiveness of the process,
will be to construct and operate a full-scale plant. Experience of South Tahoe with
other elements of the advanced waste treatment process has clearly shown that full
plant operation is the only effective means for demonstrating performance. Pilot
plant studies, while necessary and effective, do not give the proof that the process
can be operated on a continuous daily basis and that the process can work within
the limits of dependability and operating limitations imposed by full-scale plants.
Therefore, the South Tahoe Public Utility District strongly believes that the
effectiveness of the proposed ion exchange process for nitrogen removal can be
proven only by constructing and operating a full-scale TVi-mgd capacity plant.
Construction and operation of the full-scale plant will make available complete,
accurate, and dependable information on nitrogen removal and will demonstrate
that the operation can be done on a continuous, effective basis. Such information
will be of inestimable value in the coming decade, when the water quality control
program will be instigated on a nation-wide basis.
105
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It is only logical that the ion exchange plant be constructed at South Tahoe, since
the District has available the only full-scale operating advanced wastewater
treatment plant in the nation along with the trained, skilled operating personnel
required to operate the ion exchange plant, and also has available through its
manager and engineering consultants, the technical experience and expertise to make
the design, construction, and operation phases a success.
Proposed Schedule
The design of the full-scale plant can be initiated immediately upon approval being
given by FWQA. The design can be completed within a 6-month period and
construction bids can be obtained within 30 days after approval of these final plans.
The District will utilize its consulting engineers, Clair A. Hill & Associates, to
accomplish the design. The consultants will be associated with Cornell, Rowland,
Hayes and Merryfield Consulting Engineers of Corvallis, Oregon, during the design
and operation of the ion exchange plant. These consultants have previously
designed, supervised construction, and trained the operating personnel for the
District's advanced waste treatment plant.
The construction of the plant will require approximatley 9 months on an accelerated
schedule. A one-year construction period would be advisable in order not to pay a
premium for earlier completion. Construction would be done through a public
bidding procedure and the contract would be awarded to the lowest responsible
bidder. Inspection of the construction, to assure compliance with the approved plans
and specifications, would be done by the District's consulting engineers.
Allowing adequate time for review of this report, design of the final plant, and
construction of the plant, indicates that operation could begin sometime during the
spring of 1972. Since the proposed process is new and untried, an operating and data
collection program extending over a 2-year period is recommended. This would
mean that the final project report would be submitted in 1977.
106
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Project Grant Costs
The proposed project as outlined here would be financed entirely by grant funds
from the FWQA. The District would make available its advanced waste treatment
plant, laboratory facilities, and administrative offices to the project. In addition, the
District Manager, Russell L. Gulp, would serve as the Grant Project Director at no
charge to the project.
Financial requirements for the grant would therefore need to be disbursed by the
FWQA in accordance with the following schedule:
GRANT
FISCAL YEAR DESCRIPTION OF WORK AMOUNT
1970-71 Engineering Design $115,500
1970-71 Construction & Inspection 927,000
1971-72 Construction & Inspection 939,000
1972 thru 1977 Operation & Data Collection 323,000/yr
The above costs are requested to be paid by the FWQA under a continuing grant
program.
107
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o
00
VICINITY MAP
LOCATION PLAN
H-
CLAIR A. HILL &. ASSOCIATES
CONSULTING ENGINEERS
FUH.K. UTILITY
IOKI EXCMAM6E. MITeQ4CM REMOVAL PUtJT
VICIMITV MAJ» t. LOCATIOM
-------
MAJCM TYPES
r'f TREATMENT
PROVIDED
RIMARY TREATMENT
ISOLIOS SEPARATION
SECONDARY TREATMENT
(BIOLOGICAL TREATMENT)
1 BASIC
CHEMICAL TREATMENT { NITROGEN
& PHOSPHATE REMOVAL I REMOVAL
FILTRATION
ACTIVATED
CARBON
ABSORPTION
ION EXCHANGE
DISINFECTION
MARSHALL FLUMES
FLOW MEASUREMENT
AND DIVISION
SARMINUTORS
RECLAIMED
WATER TO
RESERVOIR
WASTE WATER
FLOW THRU
PLANT
SOLIDS
HANDLING LIME
AND CARBON
RECLAMATION
& ION EXCHANGE
REGENERATION
STERILE ASH TO
. DISPOSAL
CARBON
OE FINING
TANKS
CARBON TO
HE USE
ft
CLAIR A. HILL &. ASSOCIATES
CONSULTING ENGINEERS
IOVJ
UTIt-iTV DISTRICT
REMOVAL PUNT
TOTAJ- PLAKJT PLOvV
-------
INFLUENT HEADER ,
INFLUENT FEED
FROM CARBON COLUMNS
ION
EXCHANGE
BEDS __---
r
REGENEBANT
10
. r
i^?i rT-
f
] HEGENEHANT SOLUTION INLET HSABIB I I I I
1
>
WASTE
'HEGENEHANT SOLUTION
INLET HEADER
HEGENERANT
PUMPS
BACKWASH SUPPLY LINE
HEGENEHANT SOLUTION SUCTION LINE
BACKWASH
TOWER INFLUENT
PUMPS
REGENEHANT STORAGE TANKS
AMMONIA
STRIPPING
TOWER
TOWER INFLUENT LINE
TOWER EFFLUENT LINE
TOWER
STORAGE
TANK
-TOWER
RECYCLE
LINE
1 BEDS 1. 2 & 3 ARE SHOWN IN REGENERATION CYCLE
2 BEDS 4 - 12 ARE SHOWN IN SERVICE CYCLE
3 TOWER SHOWN STRIPPING FROM REGENERANT
TANK A SINGLE CYCLE
TANKS B &. C ARE IN REGERATION CYCLE.
4 ® INDICATES VALVE CLOSED
X INDICATES VALVE OPEN 4
H-
CLAIH A.
IOM EXCMA.M6E NITRO^EM
2
-------
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4" i" r
DEDUCED r LAIS'
H-
CLAIR A. HILL & ASSOCIATES
CONSULTING ENGINEERS
,\i ?EUOVAL PLAWT
VI£.C^:A.M:CA~- PI_AUS
-------
J-*
to
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CONSULTING ENGINEERS
scrtou& e.
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-------
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5
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-------
_L
Access/on Number
W
5
2
Subject Field & Group
05D
SELECTED WATER RiSOURCIS ABSTRACTS
INPUT TRANSACTION FORM
Or&anization
South Tahoe Public Utility District, South Lake Tahoe, California 95705
Title
WASTEWATER AMMONIA REMOVAL BY ION EXCHANGE
10
Aathorfe)
16
Project Designation
EPA Project #17010 ECZ
21
Note
22
Citation
23
Descriptors (Starred First)
*Waste Treatment, *Ammonia, *Ion Exchange, ^Electrolysis, *Cost Comparisons,
Sewage, Sewage Effluents, Municipal Wastes, Waste Water (Pollution), Nitrogen
Compounds, Cation Exchange, Zeolites, Separation Techniques, Air-Water
Interfaces, Cost Analysis, Design, Design Criteria
25
Identifiers (Starred First)
*Clinoptilolite, ^Stripping, Air Stripping
27
Abstract
Pilot plant investigations were conducted on the ion exchange removal of ammonia-
nitrogen from clarified and carbon-treated secondary effluents and from clarified
raw sewage. The ion exchange process utilized clinoptilolite, a natural zeolite.
Average ammonia removals from low magnesium wastewaters were in the range of 93%
to 97%. With a wastewater Mg concentration of 20 mg/1, solids formation presented
problems but they appear surmountable. The primary method used for regenerant
renovation was air stripping with which a 2N_ regenerant at a pH of 11.5 is recom-
mended. Electrolytic regenerant renovation using a neutral solution that is less
prone to solids formation was also piloted during the project.
Two process designs are included giving cost estimates for ion exchange ammonia
removal from tertiary effluent. With capital costs amortized at 6% for 20 years,
the total cost to remove ammonia from 1000 gal. of tertiary effluent is 14.8£ for a
7.5 mgd plant using regenerant air stripping and 12. 7£ for a 10 mgd plant using
electrolytic regenerant renovation. The 7.5 mgd design was prepared by South Tahoe
Public Utility District under EPA Project Number 17010 EEZ and is included for
convenience. Other work discussed in the report was performed by Battelle-
Northwest under EPA Project Number 17010 ECZ. (Mercer - Battelle-Northwest)
Abstractor
_ _ Basil
W. Mercpr
IriNtilution
Battelle-Northwest
WFi:ID2 (REV. JULY 1969)
WRSI C
SEND, WITH COPY OF DOCUMENT.
- WATER RESOURCES SCIENTIFIC INFORMATION CENTER
U.S. DEPARTMENT OF THE INTERIOR
WASHINGTON. D. C. 20240
4U.S. GOVERNMENT PRINTING OFFICE: 1972 484-483/85 1-3
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