EPA-600/2-77-146
September 1977
Environmental Protection Technology Series
WASTEWATER DEMORALIZATION BY
TWO-STAGE FIXED-BED ION EXCHANGE PROCESS
Municipal Environmental Research Laboratory
Office of Research and Development
U.S. Environmental Protection Agency
Cincinnati. Ohio 45268
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RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S. Environmental
Protection Agency, have been grouped into nine series. These nine broad cate-
gories were established to facilitate further development and application of en-
vironmental technology. Elimination of traditional grouping was consciously
planned to foster technology transfer and a maximum interface in related fields.
The nine series are:
1. Environmental Health Effects Research
2. Environmental Protection Technology
3. Ecological Research
4. Environmental Monitoring
5. Socioeconomic Environmental Studies
6. Scientific and Technical Assessment Reports (STAR)
7 Interagency Energy-Environment Research and Development
8. "Special" Reports
9 Miscellaneous Reports
This report has been assigned to the ENVIRONMENTAL PROTECTION TECH-
NOLOGY series. This series describes research performed to develop and dem-
onstrate instrumentation, equipment, and methodology to repair or prevent en-
vironmental degradation from point and non-point sources of pollution. This work
provides the new or improved technology required for the control and treatment
of pollution sources to meet environmental quality standards.
This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.
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EPA-600/2-77-146
September 1977
WASTEWATER DEMORALIZATION BY TWO-STAGE FIXED-BED
ION EXCHANGE PROCESS
by
Ching-lin Chen and Robert P. Miele
County Sanitation Districts of Los Angeles County
Whittier, California 90607
Contract No. 14-12-150
Project Officer
Irwin J. Kugelman
Wastewater Research Division
Municipal Environmental Research Laboratory
Cincinnati, Ohio 45268
MUNICIPAL ENVIRONMENTAL RESEARCH LABORATORY
OFFICE OF RESEARCH AND DEVELOPMENT
U.S. ENVIRONMENTAL PROTECTION AGENCY
CINCINNATI, OHIO 45268
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DISCLAIMER
This report has been reviewed by the Municipal Environmental
Research Laboratory, U.S. Environmental Protection Agency, and
approved for publication. Approval does not signify that the
contents necessarily reflect the views and policies of the U.S.
Environmental Protection Agency, nor does mention of trade names
or commercial products constitute endorsement or recommendation
for use.
ii
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FOREWORD
The Environmental Protection Agency was created because of
increasing public and government concern about the dangers of
pollution to the health and welfare of the American people.
Noxious air, foul water, and spoiled land are tragic testimony
to the deterioration of our natural environment. The complexity
of the environment and the interplay between its components
require a concentrated and integrated attack on the problem.
Research and development is that necessary first step in problem
solution and it involves defining the problem, measuring its
impact, and searching for solutions. The Municipal Environmental
Research Laboratory develops new and improved technology and
systems for the prevention, treatment, and management of waste-
water and solid and hazardous waste pollutant discharges from
municipal and community sources, for the preservation and treat-
ment of public drinking water supplies, and to minimize the
adverse economic, social, health, and aesthetic effects of pollu-
tion. This publications is one of the products of that research;
a most vital communications link between the research and the
user community.
Renovation of wastewater to allow for reuse often requires that
salts which are added during use be removed. This report
summarizes studies on the use of ion exchange for demineraliza-
tion of effluents from wastewater treatment plants.
Francis T. Mayo, Director
Municipal Environmental
Research Laboratory
m
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ABSTRACT
A 9.5 1/min (2.5 gpm) two-stage fixed-bed ion exchange pi-
lot plant was used to evaluate the ion exchange process for de-
mineralization of carbon-treated secondary effluent at the
Pomona Advanced Wastewater Treatment Research Facility, Pomona,
California. The pilot plant consisted of four exchange columns
operated in series — primary cation, primary anion, secondary
cation, and secondary anion column. The regenerant solutions
were applied in a downflow direction, first to the secondary
columns and then to the primary columns. This mode of regen-
eration was to ensure that the secondary exchange columns were
at a higher state of regeneration and thus were more effective
in removing the monovalent ions with lower selectivity order.
After the evaluation of the two-stage ion exchange process,
a fully automated 15 1/min (4 gpm) single-stage ion exchange pi-
lot plant was then operated over a period of thirty-two months
for resin life study. The pilot plant simulated a full scale
operation and regeneration of the primary stage of a two-stage
fixed-bed ion exchange plant. Although the physical appearance
of the resin particles was noticeably changed, yet no signifi-
cant deterioration in the overall process performance was ob-
served during the thirty-two month study period. Therefore,
the resins were considered rather stable for the demineraliza-
tion of a carbon-treated secondary effluent in a fixed-bed mode
of operation.
The estimated process cost (based on August, 1973 material
and construction costs) to demineralize the Pomona wastewater
from 600 mg/1 TDS to 60 mg/1 TDS in a 37,850 cu m/day (10 MGD)
plant was about 5.9£/l,000 liters (22.4^/1,000 gallons), ex-
cluding the costs of carbon pretreatment and brine disposal.
Cost estimates were also made for higher influent TDS waste-
water, as high as 1,500 mg/1. A reduction in ion exchange
total process cost was shown feasible by blending practice.
This report was submitted by County Sanitation Districts of
Los Angeles County in fulfillment of Contract No. 14-12-150
under the partial sponsorship of the U.S. Environmental Protection
Agency. Work on this report was completed as of June 1973.
iv
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CONTENTS
Foreword i i i
Abstract iv
Figures vi
Tables viii
Acknowledgments x
1. Introduction 1
2. Conclusions 3
3. Recommendations 5
4. Pilot Plant Description 6
5. Pilot Plant Results and Discussions 15
Pretreatment 15
Regeneration efficiency vs. regenerant level . 16
Regeneration level vs. product quality .... 16
Demineral ization flow rate 19
Carbon dioxide removal 21
Weak acid cation exchange resin 23
Effect of influent TDS 30
Brine characteristics 39
Brine recycling 39
Resin life 50
Steady-state pilot plant operation 60
6. Cost Estimate 63
References 71
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FIGURES
Number Page
1 Schematic flow diagram for demineralization cycle . . 10
2 Schematic flow diagram for backwash cycle 11
3 Schematic flow diagram for regeneration and
rinsing cycles 13
4 Regeneration study—cation exchange resin 17
5 Regeneration study-- anion exchange resin 18
6 Effect of demineralization flow rate upon resin
breakthrough capacity 20
7 Schematic regeneration flow diagfam for weak acid
cation exchange resin study 25
8 Schematic demineralization flow diagram for weak
acid cation exchange resin study 26
9 Effects of influent IDS upon IDS removal and
product IDS 38
10 Relationship of concentration vs. volume of the
impurity ions in the cation exchanger brine ... 43
11 Relationship of concentration vs. volume of the
impurity ions in the anion exchanger brine. ... 44
12 Zoning diagram of brine recycling 45
13 Schematic flow diagram for ion exchange process
without brine recycling 46
14 Schematic flow diagram for ion exchange process
with brine recycling 47
15 General layout of the automated single-stage
ion exchange pilot plant 56
vi
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FIGURES(CONTINUED)
Number Page
16 System performance of the single-stage ion
exchange unit 57
17 Performance of the two-stage ion exchange system
under optimum operating conditions 61
18 Effect of plant size on ion exchange process
cost—with 90% IDS reduction 64
19 Effect of influent IDS upon ion exchange process
cost—with 90% IDS reduction 66
vii
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TABLES
Number pag<
1 Dimensions and Specifications of Pilot Plant
Ion Exchange Resin Columns 7
2 Resin Properties and Characteristics 9
3 Typical Operating Flow Rates for Two-Stage Ion
Exchange Pilot Plant . . . *. 14
4 Effects of Carbon Dioxide Removal upon the
System Performance 22
5 Water Quality Characteristics in Phase One Weak
Acid Cation Exchange Resin Study 27
6 Comparison of the Performance of the Two-Stage
Ion Exchange System between Two Different
Modes of Operations (Weak Acid Resin
Study — Phase One) 29
7 Water Quality Characteristics in Phase Two Weak
Acid Cation Exchange Resin Study 31
8 Comparison of the Performance of the Two-Stage
Ion Exchange System between Two Different
Modes of Operations (Weak Acid Resin Study--
Phase Two) 32
9 Performance of the Two-Stage Ion Exchange
System Operating on 610 mg/1 TDS Feed Water. ... 34
10 Performance of the Two-Stage Ion Exchange
System Operating on 1,150 mg/1 TDS Feed Water. . . 35
11 Performance of the Two-Stage Ion Exchange System
Operating on 1,640 mg/1 TDS Feed Water 36
12 Effects of the Influent TDS upon the Two-Stage
Ion Exchange System Performance 37
vi 11
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TABLES(CONTINUED)
Number Page
13 Charactertistics of Brines from the Cation and
Anion Exchange Columns 40
14 Minimum Reject Streams Per Cycle of Regeneration ... 41
15 Average Effluent Qualities of the Two-Stage Ion
Exchange System Operating with Brine
Recycling 48
16 Effects of Brine Recycling upon the Two-Stage
Ion Exchange Pilot Plant Performance 49
17 Quality Characteristics of the Various Portions
of the Brines 51
18 Dry Screen Analysis of the Cation and Anion
Exchange Resins (% Retained Cumulative) 52
19 Resin Capacities of the Cation and Anion Resins. ... 53
20 Rinse Water Requirements of Anion Resins 54
21 Operation Statistics of Single-Stage vs.
Two-Stage Ion Exchange System 58
22 Typical Performance of the Automated Single-Stage
Ion Exchange System 59
23 Average Water Quality Characteristics of the
Two-Stage Ion Exchange Pilot Plant under
Optimum Operating conditions 62
24 Cost Estimate of Two-Stage Ion Exchange Process--
37,850 cu m/day (10 MGD) Product Water Plant ... 65
25 Cost Estimate Comparison of Two-Stage vs.
Single-Stage Ion Exchange System--
37,850 cu m/day (10 MGD) Plant 68
26 Cost Estimate Comparison of Two-Stage vs.
Single-Stage Ion Exchange System with Blending
Operation 69
27 Cost Estimate for Producing 500 mg/1 TDS Blended
Product Water--37,850 cu m/day (10 MGD) Ion
Exchange Plant Size 70
ix
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ACKNOWLEDGMENTS
*
This study was jointly sponsored by the U.S. Environmental
Protection Agency and the County Sanitation Districts of
Los Angeles County.
The authors are deeply grateful to Mr. Eugene P. Young of
Infilco, Inc., Tucson, Arizona, for his advice and cooperation
in this effort.
Thanks are also extended to Dr. Hans Krock, former project
engineer of the County Sanitation Districts of Los Angeles
County, for his participation in the initial stages of the pi-
lot plant study.
The valuable assistance of both the laboratory and the pi-
lot plant operating personnel of the Pomona Advanced Wastewater
Research Facility are also gratefully appreciated.
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SECTION 1
INTRODUCTION
Almost all uses of water serve to increase the mineral and
organic contents of the water- In most cases, the amount of
minerals added during one domestic use cycle of water is about
300 mg/1. If the natural quality of any water body is to be
well preserved, it is necessary that all the added impurities,
both organic and inorganic, be removed from the wastewater
streams. The organic impurities can be effectively removed by
many processes, such as biological oxidation and physical ad-
sorption, while the inorganic minerals can be most effectively
removed by demineralization processes. Therefore, wastewater
demineralization has become an indispensable part of the total
effort to conserve the natural water environment.
Furthermore, as the human activity continues to acceler-
ate, it becomes more difficult to supply a sufficient quantity
of the required quality of water. Because of its availability
in quantity and in the needed location, wastewater reuse has
emerged as a potential solution to water supply problems. The
water quality requirements for different types of wastewater
reuse are quite different. Wastewater demineralization has be-
come an effective means to reduce the TDS level to meet the
quality requirement for any particular water reuse application.
The ion exchange process, which has been successfully em-
ployed for many years in municipal and industrial water treat-
ment, has only recently been considered for demineralization
of wastewater. This delay has been primarily attributed to
the late development of the macroreticular resins, which have
much larger interstices to minimize the resin fouling caused
by the refractory organic materials present in the wastewater.
In addition to this resin development, an alteration of re-
generation concepts has contributed to the success of this
application to wastewater demineralization. In industrial
applications, resins are regenerated with excessively high
levels of regenerant to achieve complete regeneration of the
resins yielding maximum resin capacity and best effluent
quality. However, in the wastewater demineralization appli-
cations, a certain amount of TDS in the product water is
quite acceptable. For this reason, the regenerant level for
the resins can be selected as low as possible to optimize the
usage of regenerant. This has greatly improved the regenerant
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utilization efficiency from 30 percent or less in industrial
applications to 85 percent or more in wastewater applications.
Since the regenerant cost is the dominant item in the ion ex-
change unit process cost, the great improvement in the regen-
eration efficiency has substantially reduced the total process
cost. Further savings can be realized by blending of the de-
mineralized water with the non-demineralized water to produce
the desired water quality for a particular wastewater reuse
application.
A 9.5 1/min (2.5 gpm) two-stage and a fully automated 15
1/min (4 gpm) single-stage fixed-bed ion exchange pilot plant
were thus operated at Pomona Advanced Wastewater Treatment
Research Facility to achieve the following specific objectives
A. To evaluate the effect of carbon adsorption
pretreatment on the ion exchange process
performance.
B. To optimize the various operating parameters
of a two-stage fixed-bed ion exchange process
for wastewater demineralization.
C. To determine the feasibility of partial reuse
of the brine waste.
D. To investigate the resin stability and process
reliability through a long term routine
operation.
E. To develop the process cost for wastewater
demineralization by a two-stage fixed-bed ion
exchange process.
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SECTION 2
CONCLUSIONS
The following conclusions can be drawn from this pilot
plant study:
A. Activated carbon adsorption was an effective pretreat-
ment process in preventing the resins being fouled by the or-
ganic substances in the feed water.
B. The two-stage fixed-bed ion exchange pilot plant could
effectively and reliably reduce the influent TDS by 90 percent
at the regenerant levels of 17.6 g/1 (1-1 lb/cu ft) of I^SO^ and
9.6 g/1 (0.6 lb/cu ft) of NH3 for cation and anion exchanger,
respectively -
C. At the above regenerant levels, the pilot plant pro-
duced regeneration efficiencies of 85 percent and 90 percent
for the cation and anion exchanger, respectively.
D. Operation at these low regenerant levels resulted in
the cation resins being somewhat more flow sensitive than ex-
pected. The anion resins were less sensitive to flow rate with-
in the test range.
E. Partial recycling of the brines, which consisted of
spent regenerants and rinse waters, could improve the water re-
covery from 89 percent to 93 percent without degrading the
product quality.
F. The removal of C02 from the primary cation column efflu-
ent did not significantly improve the performance of the anion
exchanger during the demineralization cycle. However, it did
minimize the minor foaming problem which developed in the re-
generation cycle.
G. The combination of a weak acid primary cation exchanger
strong acid secondary cation exchanger did not produce a
r performance than the strong acid cation exchanger in both
rv and secondary staaes.
and a strong acid secondary c
better performance than the s
primary and secondary stages.
H. Similar process performance could be attained by apply-
ing the same operating conditions to the carbon-treated
secondary effluent with TDS as high as 1,640 mg/1.
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I. The results of the resin life study, which was conducted
with an automated single-stage ion exchange pilot plant, indi-
cated that both cation and anion exchange resins used in this
study were very stable.
i
J. The estimated process cost for a 90 percent deminerali-
zation of the Pomona carbon-treated secondary effluent in a
37,850 cu m/day (10 MGD) plant was about 5.94/1,000 liters
(22.44/1,000 gallons), excluding the costs of carbon adsorption
pretreatment and brine disposal. The estimate was based on
August, 1973 construction and material costs.
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SECTION 3
RECOMMENDATIONS
The following subjects which may eventually lead to a sub-
stantial reduction in ion exchange process cost are recommended
for further studies:
A. To study the possibility of improving the resin
regeneration efficiency by adopting a true countercurrent ion
exchange pilot plant system instead of the semi-countercurrent
two-stage fixed-bed system.
B. To determine the potential savings in resin inventory
in a continuous (or pulsing) type moving-bed ion exchange sys-
tem as compared to the two-stage fixed-bed pilot plant system.
C. To investigate
generant, such as lime
resin regeneraton.
the feasibility of using a
slurry instead of ammonium
low cost re-
hydroxide, for
D. To demonstrate the profitability of using a new
-NHttOH instead of HaSOit-NHitOH regeneration mode to recover
the ammonium ion from the brine in a valuable NH4N03 form.
E. To evaluate the various novel ion exchange processes,
such as Desal process (1) by Rohm & Haas, Philadelphia,
Pennsylvania and Sirotherm process (2) by ICI Australia Limited,
Melbourne, Australia, for wastewater demineralization applica-
tion.
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SECTION 4
PILOT PLANT DESCRIPTION
GENERAL
The ion exchange pilot plant, which was supplied by
Infilco, Inc., Tucson, Arizona, was installed at the Pomona
Research facility in early 1967. The ion exchange portion of
the pilot plant consisted of four columns in series — primary
cation, primary anion, secondary cation, and secondary anion.
The dimension and resin volume of each ion exchange column are
listed in Table 1. Pretreatment facilities such as densator (a
sludge blanket clarifier), rapid sand filter, and activated car-
bon adsorption column were also supplied with the original pilot
pi ant.
Since the pilot plant was not automated and was operated
for research purposes only, operational procedures as would ex-
ist in a full scale plant were not employed. Instead, operation
of the unit was dictated by the fact that personnel were at the
site only during the daytime working shift. This limitation
necessitated tailoring items such as flow rate and regeneration
interval to accommodate the operating schedule. The regenera-
tion procedure required attendance of technicians while the de-
mineralization cycle did not. Therefore, the unit was placed
on-stream to commence the demineralization cycle at 4 PM each
afternoon and this phase continued through 8 AM the following
morning. The flow rate during the demineralization cycle was
regulated so that exhaustion would not occur prior to 8 AM.
Backwashing, regeneration, and rinsing procedures were then com-
pleted within two hours and the unit remained off-stream until
the next demineralization cycle began at 4 PM. Product water
volume from each cycle of operation of the pilot plant was
approximately 9,080 liters (2,400 gallons). The same unit could
be operated at higher flow rates and regenerated more often to
allow treatment of about 18,900 liters (5,000 gallons) during
each 24 hour period.
RESIN.S
The resins used in this study were the Duolite C-20 strong
acid cation exchange resin and the Duolite A-30B intermediate
base anion exchange resin. Both resins were manufactured and
supplied by the Diamond Shamrock Chemical Company, Redwood City,
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•TABLE 1. DIMENSIONS AND SPECIFICATIONS OF PILOT PLANT ION EXCHANGE RESIN COLUMNS
Parameter
Primary
Cation
Col umn
Primary
An ion
Col umn
Secondary
Cation
Col umn
Secondary
An ion
Col umn
Column Diameter
in
cm
Column Height
in
Resin
Resin
Type
cm
Bed Depth
in
cm
Vol ume
cu ft
cu m
of Resin
20
50.8
72
183
45
113
8.0
0.23
Duol ite
C-20
10
25.4
56
142
39
98.7
1.6
0.05
Duol i te
A-30B
13
33.0
58
147
41
105
3.2
0.09
Duol ite
C-20
10
25.4
56
142
39
98.7
1 .6
0.05
Duol ite
A-30B
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California. Some of the major physical and chemical character-
istics of the two resins are shown in Table 2. The A-30B in-
termediate base resin was regenerated with ammonium hydroxide in
this study. Therefore, it was more appropriately classified as
a weak base anion exchange resin.
OPERATION
Basically, each test run of the ion exchange pilot plant in-
volved four consecutive operation cycles, namely, demineraliza-
tion cycle, backwash cycle, regeneration cycle, and rinse cycle.
During the demineralization cycle, flow was directed down-
ward through each of the four columns in series as shown in
Figure 1. The source of the feed water was the carbon-treated
secondary effluent from the 12.6 I/sec (200 gpm) carbon adsorp-
tion pilot plant at the Pomona Research Facility. Some of the
product water from the ion exchange pilot plant was used as pro-
cess water for chemical make-up and rinse operation, and the
rest of it was discharged to waste. The demineral ization cycle
was first monitored by the determinations of pH and conductivity
in the successive grab samples of the product water. This moni-
toring operation was soon replaced by the on-stream conductivity
recording system.
Backwashing was accomplished in an upflow direction com-
mencing with the secondary anion exchanger and proceeding in re-
verse order to the primary cation exchanger. Backwash water
(carbon-treated secondary water) for each column was first
passed downflow through the preceding columns of the series as
shown in Figure 2.
Backwashing was necessary to remove suspended solids and
to rearrange the resin beds. Most of the suspended solids were
removed in the primary cation column, and backwashing of this
column was conducted after each demineralization cycle. Re-
arranging the resin bed was necessary primarily to disperse the
uppermost layer of the cation resins, which were exhausted in
the calcium form, so that the potential tendency of calcium
sulfate precipitation could be greatly reduced during regenera-
tion with sulfuric acid. Backwashing also eliminated short-
circuiting by removing entrapped bubbles or other restrictions.
A backwash frequency of once every 5 demineralization cycles
was found adequate for the secondary cation column and for both
anion columns.
Although the backwash wastewater was piped to waste in the
pilot plant, this water could be recycled to the pretreatment
unit (carbon adsorption or equivalent) in a full-scale plant.
The backwash wastewater was of no higher mineral content than
the feed water, but it might be necessary to remove the sus-
pended solids before recycling.
8
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TABLE 2. RESIN PROPERTIES AND CHARACTERISTICS
Parameter
Duolite C-20
Duolite A-30B
Amberlite IRC-84
Functional Groups
Ionic Form
Physical Form
Chemical Classification
Specific Gravity
Screen Grading (Wet) **
Effective Size, mm
Uniformity Coefficient
Total Exchange Capacity,
Nuclear Sulfonic
Hydrogen
Beads
Strong-acid
1.32 (Na form)
16-50 mesh
0.53
1.49
Tertiary Amine,
Quarternary Ammonium
Free Base
Beads
Intermediate-base *
1.18 (Cl form)
16-50 mesh
0.47
1.62
Carboxylic
Hydrogen
Beads
Weak-acid
1.16 (H form)
16-50 mesh
0.42
1.75
equivalent/liter
Effective pH Range
Maximum Temperature, °C
Moisture Content, %
Manufacturer
2.2
0-14
150
43-46
Diamond Shamrock
2.6
0-9
80
53-57
Chemical Co. (1)
4.1
5-14
120
43-50
Rohm & Haas
Co. (2)
* Being considered as a weak-base resin due to the use of ammonium hydroxide as regenerant in this
study.
** U.S. Standard Screens
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FEED
WATER
PRODUCT
WATER
Figure I. Schematic flow diagram for demineralization cycle.
10
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FEED
WATER
BACKWASH
WASTE
SEQUENCE
OF
BACKWASH
I
PRODUCT
WATER
NUMBER OF VALVE OPENED
DURING BACKWASH CYCLE
SECONDARY ANION COLUMN : © , ® , © , (H) , ©
SECONDARY CATION COLUMN '• © , ® , © , ®
PRIMARY ANION COLUMN = ©,©•(§)
PRIMARY CATION COLUMN ' © , ®
Figure 2. Schematic flow diagram for backwash cycle.
11
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Regeneration of the exhausted resin was performed in a
downflow direction, with the regenerant passing through the
secondary columns first and then through the primary columns.
The cation columns were regenerated with a 4 percent sulfuric
acid solution, while the anion columns were regenerated with a
4 percent ammonium hydroxide solution. Regeneration of the
secondary columns first resulted in a higher state of regenera-
tion for the secondary columns to ensure an effective removal of
ions which were not removed in the primary columns due to a
relatively low selectivity.
After regeneration, the excess regenerant solution should
be rinsed from the resin beds prior to commencing a deminerali-
zation cycle. The first part of the rinsing operation, a slow
rinse, was employed to force the regenerant solution from the
secondary column through the primary column to complete the re-
generation itself. The second part of the rinsing operation, a
fast rinse, was employed to remove the excess regenerants from
the primary cation and primary anion columns. The slow rinse
operation was accomplished with demineralized water, at the same
flow rate of the regenerant solution, while the fast rinse
operation was accomplished with the feed water. The fast rinse
flow rate can be much higher than the slow rinse flow rate.
However, due to the limiting capacity of the feed pump in the
pilot plant, the fast rinse flow rate for the primary cation
column was actually slower than the slow rinse flow rate. A
schematic flow diagram of the regeneration and rinsing opera-
tions is shown in Figure 3. The flow rates employed for each
type of operation are shown in Table 3.
12
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«O
FAST RINSE
CO
c
3 ^
OJ
§.£ o
-i » ^ r»=
§'i § ^
5s- m '
o.i
• Q
«Q
•^
O
3-—
^—*-
(D
(O
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TABLE 3. TYPICAL OPERATING FLOW RATES
FOR TWO-STAGE ION EXCHANGE PILOT PLANT
Operation Cycle
Demineral i zation
gpm
gpm/sq ft
gpm/cu ft
Backwash
gpm
gpm/sq ft
Regeneration
gpm
gpm/cu ft
Slow Rinse
gpm
gpm/cu ft
Fast Rinse
gpm
gpm/cu ft
Primary
Cation
2.5
1.2
0.3
5.0
2.3
5.6
0.7
5.6
0.7
5.0 *
0.6 *
Primary
An ion
2.5
4.6
1 .6
2.0
3.6
1 .6
1.0
1.6 '
1 .0
4.0
2.5
Secondary Secondary
Cation Anion
2.5 2.5
2.7 4.6
0.8 1.6
2.0 2.0
2.2 3.6
5.6 1.6
1.7 1.0
5.6 1.6
1.7 1.0
* Due to the limitation of pump capacity, the fast rinse flow
rate was slower than the slow rinse flow rate.
1 gpm = 3.8 1/min; 1 gpm/sq ft = 40.7 Ipm/sq m;
1 gpm/cu ft = 0.13 lpm/1
14
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SECTION 5
PILOT PLANT RESULTS AND DISCUSSIONS
PRETREATMENT
It is desirable to minimize.both suspended solids and or-
ganic substances in the feed water to an ion exchange system.
Suspended solids will force the ion exchange column, particular-
ly the primary cation column which is in the lead position, to
act as a filter unit. This may lead to clogging of the resin
bed and fouling of the resins. Soluble organic substances can
adversely affect resin performance, especially the anion resins,
by taking up ion exchange sites on the resins.
The pilot plant was supplied with a pretreatment unit con-
sisting of lime coagulation and sand filtration. The secondary
effluent from the Pomona Water Renovation Plant (an activated
sludge process) was treated by this pretreatment unit prior to
application on the primary cation column. In addition to the
removal of suspended solids, the pretreatment unit was also de-
signed for reducing the hardness and phosphate in the feed water
by lime precipitation. However, the experimental results indi-
cated that there was no significant savings in overall process
cost for demineralizing Pomona secondary effluent by removing
some hardness and phosphate prior to ion exchange process. Con-
sequently, the original pretreatment unit was bypassed and a
carbon adsorption system was used to pretreat the feed water as
discussed in the following paragraph.
Initially, the removal of the soluble organic substances
was to be accomplished by a carbon adsorption column located
between the primary cation column and the primary anion column.
This arrangement was to take advantage of the low pH value in
the effluent from the primary cation column. Carbon adsorption
is somewhat more efficient in treating water with a low pH
value. However, considerable problems were encountered with
leaching of iron and other metallic elements from the carbon
with subsequent precipitation of these minerals in the primary
anion column. Therefore, satisfactory operation of the carbon
column at this location was never achieved. This situation,
coupled with the fact that the chemical coagulation unit was
not achieving a substantial reduction in hardness, prompted a
15
-------
revision of the flow pattern after three months of operation.
The chemical coagulation, sand filtration, and carbon adsorption
units which were supplied with the ion exchange pilot plant were
all bypassed, and the effluent from a 12.6 I/sec (200 gpm) car-
bon adsorption pilot plant at the same site was applied directly
to the primary cation column. All data reported in this report
were based on the revised pretreatment scheme.
REGENERATION EFFICIENCY vs. REGENERANT LEVEL
Since the regeneration process involves the application of
expensive chemical regenerants, the efficiency of regeneration
is an overwhelmingly important factor in process cost considera-
tions. The regeneration efficiency and the regenerant level are
defined in this report as follows:
Regeneration efficiency (%)
_ Total equivalents of ions removed from the feed water x ,QQ
Total equivalents of regenerant applied
Regenerant level (g/1 or Ib/cu ft)
= Total weight of pure regenerant applied per unit volume of
resins regenerated
A series of experiments was conducted to determine the re-
lationship between the regenerant level and the regeneration
efficiency for both cation and anion resins. The results of
these tests are presented in Figure 4 and Figure 5 for the
cation and anion exchange resins, respectively.
As indicated in both Figure 4 and Figure 5, the regenera-
tion efficiency increased linearly as the regenerant level de-
creased for both the cation and anion resins over the range in-
vestigated. Since the regenerant level directly affects the
product water quality, the regenerant level to be applied to a
process cannot be selected simply on the basis of the results of
the regeneration efficiency study. The relationship between the
regenerant level and the product water quality was thus thor-
oughly investigated during the pilot plant study.
REGENERANT LEVEL vs. PRODUCT QUALITY
The product quality of the ion exchange process is indi-
cated in this study by the average percentage of ion leakage in
the product water. The average percentage of ion leakage for
either cation or anion exchanger is defined as the ratio of the
total equivalents of cations or anions remaining in the product
water to the total equivalents of cations or anions in the feed
water.
16
-------
100
80
z
UJ
o
UJ
cc
111
uJ
CD
UJ
tr
60
40
20
20
0,5
A
I
1.0 1.5 2.0
REGENERANT LEVEL, Ibs H2S04/cu ft
2.5
16
10
12
8
CD
111
UJ
CD
UJ
0
3.0
Figure 4. Regeneration study-cation exchange resin.
-------
88
UJ
o
U.
U-
HI
100
80
UJ
z
UJ
CD
UJ
DC
60
40
20
A
1
(Ilb/cuft=l6g/l)
I
I
0.5 1.0 1.5
REGENERANT LEVEL, Ibs NH3/cuft
25
20
UJ
15 2
o
UJ
10 3
DC
UJ
0
2.0
Figure 5. Regeneration study-anion exchange resin.
-------
The results of this series of study are included in Figure
4 and Figure 5 for cation and anion resins, respectively. As
clearly shown in these figures, both cation and anion leakage
curves show a definite transition point where leakage levels
off. The ion leakage is rapidly increased as the regenerant
level decreases below a particular level. The leakage at this
transition point is about 6 percent as indicated in both figures
for cation and anion resins. The respective regenerant levels
for cation and anion exchange resins at their transition points
are about 17.6 g/1 (1.1 Ib/cu ft) of H2S04 and 9.6 g/1 (0.6
Ib/cu ft) of NHs. The regeneration levels correspond to re-
generation efficiencies of 85 percent and 90 percent for cation
and anion exchange resins, respectively.
Most ion exchange applications require that product water
leakage be very low and thus demand a high regenerant level.
However, this is not the case for demineralization of waste-
water- The fact that ion leakages at the 6 percent level are
perfectly acceptable has allowed the use of low regenerant
levels to attain high efficiencies of regenerant utilization
and thus has made the ion exchange process an economically
promising method for wastewater demineralization.
DEMORALIZATION FLOW RATE
The capital cost of an ion exchange system is greatly
affected by the flow rate of feed water through the ion exchange
columns in the demineral ization cycle. If the demineralization
flow rate, expressed as lpm/1 (or gpm/cu ft), can be increased
without degrading the product water quality, the required resin
volume can be decreased.
To establish the maximum demineral ization flow rate, tests
were conducted for both the anion and cation resins. The cation
resin experiment was accomplished using the small secondary
cation column, instead of the large primary cation column, to
allow a wider range of flow rate for the experiment with the ex-
isting feed pump. The effect of flow rate upon the anion resin
was determined by using the primary anion column preceded by the
primary cation column to simulate the full scale operation. The
regenerant levels used for this flow rate study were about 17.6
g/1 (1.1 Ib/cu ft) of H2SOi, and 12.8 g/1 (0.8 Ib/cu ft) of NH3
for cation and anion exchange resins, respectively. The flow
rate was varied and the breakthrough capacity, expressed in
equivalents, was determined for each flow rate. Breakthrough
capacity was defined as the sum of the equivalents of ions re-
moved within a demineralIzation cycle which was ended with a
steep rise in the effluent cation or anion concentrations. The
results of these tests are shown in Figure 6.
19
-------
INi
O
o
>
'5
o-
O
o
o
QL
I
LU
jr
CD
30
25
20
i i r
/ANION EXCHANGER
O—* o——
CATION EXCHANGER
(1 gpm/cu ft=O.I3 Ipm/l)
1
1
I
O
0.5
1.0 1.5 2.0
FLOW RATE, gpm/cu ft
2.5
3.0
Figure 6. Effect of demineralization flow rate upon resin
breakthrough capacity.
-------
As shown in Figure 6, the cation breakthrough capacity de-
creased by 13 percent as the flow rate was increased from 0.13
lpm/1 to 0.33 lpm/1 (1.0 gpm/cu ft to 2.5 gpm/cu ft). The re-
duction in anion breakthrough capacity was shown to be only
about 2 percent with the same range of flow rate increase from
0.13 lpm/1 to 0.33 lpm/1 (1.0 gpm/cu ft to 2.5 gpm/cu ft). The
figure also indicates that the maximum flow rates for retaining
the maximum breakthrough capacities at the selected regenerant
levels are approximately 0.2 lpm/1 (1.5 gpm/cu ft) and 0.26
lpm/1 (2.0 gpm/cu ft) for cation and anion exchange resins, re-
spectively.
The fact that the cation resin was more flow sensitive than
the anion resin within the tested range was somewhat unexpected
based on the technical, data supplied by the resin manufactur-
er^). However, the technical data were developed at consider-
ably higher regenerant levels with better effluent water quali-
ties. The substantially low regenerant level employed at this
study was apparently responsible for the flow rate sensitivity
displayed by the cation resin.
CARBON DIOXIDE REMOVAL
A series of investigations was conducted to determine the
effect of carbon dioxide removal upon the overall performance
of the ion exchange system. The carbon dioxide was removed by
air stripping in an aeration tower installed between the pri-
mary cation and primary anion columns.
During this phase of the study, the flow rate through the
ion exchange system was maintained at 9.5 1/min (2.5 gpm). The
regenerant levels for the cation and anion resins were main-
tained at 17.6 g/1 (1.1 Ib/cu ft) of ^SOi* and 9.6 g/1 (0.6
Ib/cu ft) of NHs, respectively. Under these operating condi-
tions, the stripping tower was able to remove 90 percent of the
carbon dioxide from the primary cation column effluent, which
usually had a pH of about 2.7. The results of these experi-
mental runs are summarized in Table 4. The column A and column
B on Table 4 represent, respectively, the average of experi-
mental data of the operations without and with carbon dioxide
removal in the process. The regeneration conditions and flow
rates were identical for both modes of operations.
As indicated in Table 4, the TDS reduction in both opera-
tions were about the same. However, the removal efficiencies
for cations were either the same or significantly increased,
while the removal efficiencies for anions were all slightly de-
creased as a result of 90 percent carbon dioxide removal from
the primary cation column effluent. The carbon dioxide removal
also slightly reduced the regeneration efficiency for the anion
exchanger.
21
-------
TABLE 4. EFFECTS OF CARBON DIOXIDE REMOVAL
UPON THE SYSTEM PERFORMANCE
Parameter
Cal ci urn
Magnesi ura
Sodi um
Potassi um
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as POiJ
pH
Conducti vity (ymho s/ cm) 1
Alkalinity (as CaC03)
TDS
Feed Water
mg/1
A B
53
17
126
14
20
72
135
2.9
27
7.4
,040
218
610
62
14
107
12
22
65
113
1 .2
26
7.6
980
275
540
Product Water
mg/1
A B
0.6
0.0
15
1 .9
3.8
1 .3
14
0.35
0.25
5.8
100
40
70
0.9
0.0
7.8
1 .3
1 .5
3.3
14
0.27
0.41
6.8
60
15
65
Removal
%
A B
99 99
100 100
88 93
86 89
81 93
98 95
90 88
88 78
99 98
89 88
Column A : without C02 removal.
Column B : with COz removal.
22
-------
Although the high degree of carbon dioxide removal was not
found to improve the overall performance of the ion exchange
system during its demineralization cycle, the carbon dioxide re-
moval did solve some minor foaming problems involved in the re-
generation operation when carbon dioxide removal was not prac-
ti ced.
At the end of the above carbon dioxide removal study, the
location of the aeration tower was moved down to the end of the
ion exchange system. The product water from the ion exchange
system was pumped through the aeration tower in an attempt to
raise the pH of the product water by the removal of carbon di-
oxide from the product water. The results showed that the
aeration tower could raise the pH of the product water from pH
5.9 to pH 6.8.
WEAK ACID CATION EXCHANGE RESIN
Since a weak acid cation exchange resin could be regener-
ated more efficiently than a strong acid cation exchange resin
with a subsequent savings in acid consumption cost, it was of
great interest to investigate the applicability of a weak acid
cation exchange resin to the wastewater demineralization sys-
tem. The particular weak acid cation exchange resin selected
for this study was the Rohm & Haas IRC-84 resin. This resin
contained carboxylic acid functional groups whereas the strong
acid cation exchange resin, C-20. contained sulfonic acid func-
tional groups.
According to the following reaction equations, the strong
acid cation exchange resin, represented as R-S03H, is capable
of removing cation associated with any anion by an equivalent
exchange for hydrogen ion. However, the weak acid cation ex-
change resin, represented as R-COOH, can only effectively re-
move cation associated with bicarbonate ion.
A. Strong acid cation exchange resin:
R-S03H + NaCl •=-— R-S03Na + HC1
2R-S03H + Ca (HC03)2-^(R-S03 )2 Ca + 2H20 + 2C02
B. Weak acid cation exchange resin:
R-COOH + NaCL 5?=— R-COONa + HC1
2R-COOH + Ca (HC03)2~^=(R-COO)2 Ca + 2 H20 + 2C02
Because of the limited capability of the weak acid cation
exchange resin in cation removal, a combination of weak acid
23
-------
cation exchange resin and strong acid cation exchange resin was
employed in this study to remove all forms of cations as com-
pletely as possible. The original strong acid cation exchange
resin, C-20, in the primary cation column was replaced with a
new IRC-84 weak acid cation exchange resin, while the original
C-20 strong acid cation exchange resin in the secondary cation
column was retained for the combined system. The resin volume
in the primary cation column was reduced from its previous 0.23
cu m (8 cu ft) to 0.17 cu m (6 cu ft) to provide adequate free-
board in the column for resin bed expansion during backwashing.
However, the total amount of sulfuric acid used for the regen-
eration of this combined system was maintained at the same level
as used in previous straight strong acid cation exchange resin
system, which resulted in a slight increase of the regenerant
level from 17.6 g/1 (1.1 Ib/cu ft) to 20.8 g/1 (1.3 Ib/cu ft)
of H2SOIt for the combined cation exchange system.
The flow patterns and concentrations of regenerants for the
regeneration of both cation and anion exchange resin columns are
shown in Figure 7. As indicated in the figure, the 4 percent
sulfuric acid solution was first pumped through the secondary
cation column to regenerate the strong acid cation exchange
resin. The partially spent regenerant from the secondary cation
column was collected in a dilution tank to reduce the concentra-
tion down to the level of 0.5 percent. This diluted acid was
then pumped through the primary cation column for the regenera-
tion of the weak acid cation exchange resin. This dilution step
was necessary to prevent precipitation of calcium sulfate within
the weak acid cation exchange resin bed.
During the course of this study, two different deminerali-
zation flow schemes were investigated. In the first flow
scheme, the primary anion exchange column (with A-30B anion ex-
change resin) remained at the second position between the pri-
mary cation exchange column (with IRC-84 cation exchange resin)
and the secondary cation exchange column (with C-20 cation ex-
change resin) as shown in Figure 8-a. In the second flow
scheme, the secondary cation exchange column was placed directly
after the primary cation exchange column as shown in Figure 8-b.
Although the demineralization flow scheme was different during
this study, identical regeneration conditions as shown in Figure
7 were employed throughout the entire series of weak acid cation
exchange resin study. The results are discussed in the follow-
ing sections with respect to different flow schemes.
Phase One (IRC-84->-A-30B-*-C-20->-A-30B Flow Scheme)
The average water quality achieved by each ion exchange
column in this flow scheme is shown in Table 5. As indicated in
the table, the weak acid cation exchange column effectively
24
-------
m
H
m
m
z
ni
o>
I
-j
0.5%
H2S04
T
L
CATION COLUMN
(IRC-84)
ANION COLUMN
(A-30B)
CATION COLUMN
(C-20)
4%
H2S04
I
ANION COLUMN
(A-30B)
4%
NH40H
DEMINERALIZED
WATER
-------
FEED
1
If
§ *
2 a
5
O
1
oo
010
1
OO
OCM
_ I
o
1
8?
Z<
2*~*
z
<
PRODUCT
(a) SCHEMATIC DEMINERALIZATION FLOW DIAGRAM IN PHASE ONE
FEED
1
o ?
0 o
z ac
o tj
1
o o
O CM
O
1
il
1
8?
o*
<
<
PRODUCT
(b) SCHEMATIC DEMINERALIZATION FLOW DIAGRAM IN PHASE TWO
Figure 8. Schematic demineralization flow diagram for weak acid
cation exchange resin study.
-------
TABLE 5. WATER QUALITY CHARACTERISTICS IN PHASE ONE
WEAK ACID CATION EXCHANGE RESIN STUDY
Parameter
(mg/l)
Calcium
Magnesium
Sodium
Potassium
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as PO^)
pH
Conductivity (ymhos/cm)
Alkalinity (as CaC03)
Acidity (as CaC03)
TDS
Carbon-
Treated
Secondary
Effluent
68
14
122
12
9.7
74
140
3.9
34
7.4
1,030
234
591
Primary
Cation
Col umn
Effluent
7.0
0.64
95
7.1
8.8
74
130
3.2
33
3.4
830
35
Primary
An ion
Col umn
Ef f 1 uent
7.6
0.61
85
7.2
8.8
2.2
167
3.6
27
4.9
620
10
Secondary
Cation
Column
Ef f 1 uent
1.3
0.20
29
2.4
3.0
2.0
165
3.7
27
2.7
1,570
179
Secondary
An ion
Column
Ef f 1 uent
1.1
0.14
29
2.4
3.1
0.80
28
1.6
3.8
5.7
178
45
98
27
-------
removed the calcium and magnesium bivalent ions. However, its
capability in removing the monovalent ions, such as sodium,
potassium, and ammonium ions, was extremely poor in this study.
The total removal of the cations by the weak acid cation ex-
changer, IRC-84, was about 5.5 equivalents which represented
about 51 percent of the total cations in the feed water. This
removal efficiency was lower than the 68 percent efficiency as
achieved by the strong acid cation exchanger, C-20, under
equivalent conditions of operation. The lower removal effi-
ciency by the weak acid cation exchanger was primarily due to
its limiting exchange capacity as effected by the concentration
of the bicarbonate ion in the feed water, which approximately
amounted to 4.7 equivalents in this study. According to the
data in Table 5, the weak acid cation exchanger was able to re-
move only about 0.8 equivalents of cations beyond the effective
limit of the 4.7 equivalents set by the bicarbonate ions in the
feed water. This additional 0.8 equivalents of cations were be-
lieved to be mostly associated with sulfate ions.
According to the ion removal mechanisms of a weak base
anion exchanger, such as the A-30B in this application, the
total amount of anion removed by the A-30B resins in the pri-
mary anion exchange column in this study would be more or less
limited to the 0.8 equivalents of sulfate ions in the form of
^SOi*. However, a close check of the sulfate ion removal by
the primary anion exchanger revealed that a total of 1.5 equiva-
lents of sulfate ions were removed as indicated in Table 5. The
unexpected additional removal between 1.5 equivalents and 0.8
equivalents might have been achieved by an exchange mechanism,
instead of adsorption, as evidenced by the almost equivalent in-
crease of the chloride ions in the primary anion column efflu-
ent. The chloride ions replaced by the sulfate ions were possi-
bly left on the anion exchanger from previous operations. This
special chloride and sulfate ion exchange phenomenon was not
observed when strong acid cation exchanger, instead of weak acid
cation exchanger, was used in the primary cation exchange
column. The reason for the special behavior is not known. How-
ever, it is believed that the exchange process between the sul-
fate ions in the water and chloride ions remaining on the anion
exchange resins could be minimized if the active ionic forms
available on the anion exchange resins would be all of free
base.
Table 6 shows the comparison of the performance of the ion
exchange pilot system between two different modes of operations.
The TDS reduction was decreased from 88 percent to 84 percent
when weak acid cation exchange resin was used to replace the
strong acid cation exchange resin in. the primary cation column.
Since the resin regeneration efficiency was defined in this
study as the ratio of the total equivalents of ions removed from
28
-------
TABLE 6. COMPARISON OF THE PERFORMANCE OF THE TWO-STAGE
ION EXCHANGE SYSTEM BETWEEN TWO DIFFERENT MODES OF
OPERATIONS (WEAK ACID RESIN STUDY - PHASE ONE)
Feed Water
mg/1
Parameter
A B
Cal cium
Magnesi urn
Sodi urn
Potassi um
Ammoni um
Sul fate
Chi oride
Nitrate
Phosphate
PH
Conductivity
AT kal inity
TDS
62
15
127
12
(as N) 18
65
131
(as N) 3.9
(as P04) 37
7.5
(ymhos/cm) 980
(as CaC03) 250
594
68
14
122
12
9.7
74
140
3.9
34
7.4
1 ,030
234
591
Product Water
mg/1
A B
1
0
9
2
1
0
7
0
0
5
70
26
70
.3
.0
.6
.4
.7
.99
.8
.59
.25
.9
1 .1
0.14
29
2.4
3.1
0.80
28
1 .6
3.8
5.7
178
45
98
Removal
%
A B
98
100
92
80
91
99
94
85
99
88
98
99
76
80
68
99
80
59
89
84
Column A : with C-20 resin in the primary cation column-
Column B : with IRC-84 resin in the primary cation column
29
-------
the feed water to the total equivalent of regenerant applied,
the calculated regeneration efficiency for the combined cation
exchanger was significantly reduced from 85 to 80 percent as a
result of the limiting capacity of the weak acid cation ex-
changer in removing total cations. The, fact that the weak acid
cation exchanger accomplished an unfavorably lower regeneration
efficiency in this study has disregarded the possible savings
in regenerant consumption to restore the exchange capacity in a
weak acid cation exchanger.
Phase Two (IRC-84-*C-20->A-30B-»-A-30B Flow Scheme)
The average water qualities of the various column efflu-
ents are shown in Table 7. Only about 25 percent of the sodium,
potassium, and ammonium ions was removed by the weak acid
cation exchanger as indicated in Table 7. This led to a high
slippage of monovalent cations from the weak acid cation ex-
changer, and thus it caused substantial increase of the cation
loading on the strong acid cation exchanger. Since the total
regenerant applied to the strong acid cation exchanger was main-
tained at the same value, the increase of the cation loading
would directly increase the ion leakage from the strong acid
cation exchanger. The results indicated that cation leakage
of 17.2 percent was obtained from this combined system, while
it was only about 6.5 percent for the previous straight strong
acid cation exchanger system. The anion leakage for this com-
bined system was also increased from previous 7.5 percent to 19
percent as a result of the unusual chloride and sulfate ion ex-
change phenomenon. The comparable overall cation and anion
leakage values for the operations in the aforementioned phase
one study were about 15 percent and 17 percent, respectively.
The TDS removal in this study was decreased from previous
88 percent to 80 percent as indicated in Table 8, which was a
larger reduction than that of the phase one study. However,
the regeneration efficiency for the overall cation exchanger
was slightly improved from 80 percent in phase one study to 82
percent in this phase two study- The anion exchanger was still
maintained at the usual 90 percent level in this study.
In summary, all the experimental results obtained from both
phases of the study on the weak acid cation exchanger substan-
tiated the conclusion that a combined weak and strong acid
cation exchange resin system was less effective than an entire
strong acid cation exchange resin system for wastewater de-
mineral izati on.
i
EFFECT OF INFLUENT TDS
Following the conclusion of the study on the weak acid cation
exchange resin, another study was initiated to investigate the
30
-------
TABLE 7. WATER QUALITY CHARACTERISTICS IN PHASE TWO
WEAK ACID CATION EXCHANGE RESIN STUDY
Parameter
(mg/D
Calcium
Magnesium
Sodium
Potassium
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as POu)
PH
Conductivity (ymhos/cm)
Alkalinity (as CaC03)
Acidity (as CaC03)
TDS
Carbon-
Treated
Secondary
Ef f 1 uent
62
13
120
13
12
77
118
2.7
36
7.6
887
231
568
Primary
Cation
Col umn
Ef f 1 uent
7.6
0.45
91
8.3
8.1
82
123
2.0
33
3.4
743
35
Secondary
Cation
Column
Ef f 1 uent
1.3
0.0
30
3.5
2.8
84
122
2.3
29
2.7
1,568
204
Primary
An ion
Col umn
Effluent
3.7
0.0
33
3.7
2.8
4.9
158
2.6
33
2.8
1,307
159
Secondary
An ion
Col umn
Effluent
2.5
0.0
31
3.4
3.5
0.63
34
1.2
3.2
5.9
170
44
112
31
-------
TABLE 8. COMPARISON OF THE PERFORMANCE OF THE TWO-STAGE
ION EXCHANGE SYSTEM BETWEEN TWO DIFFERENT MODES OF
OPERATIONS (WEAK ACID RESIN STUDY - PHASE TWO)
Feed Water
mg/1
Parameter
A B
Cal cium
Magnesi urn
Sodi um
Potassi um
Ammonium
Sulfate
Chloride
Nitrate
Phosphate
pH
Conductivity
Alkalinity
TDS
62
15
127
12
(as N) 18
65
131
(as N) 3.9
(as POiJ 37
7.5
(ymhos/cm) 980
(as CaC03) 250
594
62
13
120
13
12
77
118
2.7
36
7.6
887
231
568
Product Water
mg/1
A B
1
0
9
2
1
0
7
0
0
5
70
26
70
.3
.0
.6
.4
.7
.99
.8
.59
.25
.9
2.5
0.0
31
3.4
3.5
0.63
34
1 .2
3.2
5.9
170
44
112
Removal
%
A B
98
100
92
80
91
99
94
85
99
88
96
100
74
73
71
99
71
53
95
80
Column A : with C-20 resin in the primary cation column.
Column B : with IRC-84 resin in the primary cation column.
32
-------
effects of influent TDS upon the various performance parameters
of the two-stage ion exchange system. The weak acid cation ex-
change resin, IRC-84, was removed from the primary cation
column and replaced by the original C-20 strong acid cation
exchange resin. The flow sequence of the resin columns in the
demineralization cycle was reverted to its original pattern as
shown in Figure 1.
In addition to the normal 610 mg/1 TDS level in Pomona
wastewater, two higher TDS levels were also investigated in this
study. These higher TDS feed waters were artificially obtained
by adding the appropriate amount of chemicals into the carbon-
treated secondary effluent from Pomona activated sludge plant.
The average water characteristics of the three different TDS
feed waters, namely 610 mg/1, 1,150 mg/1, and 1,640 mg/1, are
shown in Table 9 through Table 11. The amount of each indi-
vidual ion to be added to increase the TDS in the feed waters
was determined by the average concentration of that ion in some
typical wastewaters which had the TDS values about the same
levels as used in this study-
The same regeneration flow patterns as shown in Figure 3
were employed for both cation and anion exchangers throughout
this study. The regenerant levels for the cation and anion
exchangers were maintained at 17.6 g/1 (1.1 Ib/cu ft) of H2$0it
and 9.6 g/1 (0.6 Ib/cu ft) of NH3, respectively. Since both
resin volumes and regenerant levels were not increased in
accordance with the increase of the influent TDS, the volume of
the product water for the demineralization cycle was thus re-
duced from 9,080 liters (2,400 gallons) for 610 mg/1 TDS feed
water to 4,540 liters (1,200 gallons) and 3,560 liters (940
gallons) for the higher TDS feed waters of 1,150 mg/1 and 1,640
mg/1, respectively. The demineral ization flow rate was in-
creased from 9.5 1pm (2.5 gpm) for operation with 610 mg/1 TDS
feed water to 15 1pm (4 gpm) for operations with the two higher
TDS feed waters. This increase in the feed rate would allow
the entire operation run for the higher TDS feed waters to be
easily monitored during the normal daytime working hours. From
Figure 6 and Table 3, it is evident that this higher flow rate
would not significantly affect the resin performance.
The performance of each individual column of the two-stage
ion exchange system for this study is shown, respectively, in
Table 9 through Table 11 for 610 mg/1, 1,150 mg/1, and 1,640
mg/1 feed waters. Further analyses of the experimental data are
made to demonstrate the effects of the influent TDS upon the
various performance parameters as shown in Table 12. The
effects of influent TDS upon TDS removal and the residual TDS
in the product water are also illustrated in Figure 9.
33
-------
TABLE 9. PERFORMANCE OF THE TWO-STAGE ION EXCHANGE SYSTEM OPERATING ON 610 mg/1 TDS FEED WATER
Parameter
(mg/1)
Calcium
Magnesium
Sodi urn
Potassium
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as P04)
PH
Conductivity (ymhos/cm)
TDS
Alkalinity (as CaC03)
Acidity (as CaC03)
Carbon
Col umn
Ef f 1 uent
53
17
126
14
20
72
135
2.9
27
7.4
1,040
610
218
Primary
Cation
Ef f 1 uent
2.0
0.59
61
7.3
9.6
72
132
2.8
27
2.7
1,390
298
no
Primary
An ion
Effluent
1.7
0.56
59
7.1
9.2
3.6
84
1.6
15
5.7
390
198
51
Secondary
Cation
Ef f 1 uent
1.1
0.38
16
1.9
4.0
3.6
83
1.5
14
2.8
1,040
104
110
Secondary
An ion
Ef f 1 uent
0.60
0.0
15
1.9
3,8
1.3
14
0.35
0.25
5.8
100
70
39
Total
Removal
%
99
100
88
86
81
98
90
88
99
90
89
-------
TABLE 10. PERFORMANCE OF THE TWO-STAGE ION EXCHANGE SYSTEM OPERATING ON 1,150 mg/1 TDS FEED WATER
CO
01
Parameter
(mg/1 )
Calcium
Magnesium
Sodium
Potassium
Ammonium
Sulfate
Chloride
Nitrate
Phosphate
PH
(as N)
(as N)
(as POJ
Conductivity (ymhos/cm)
TDS
Alkalinity
Acidity
(as CaC03)
(as CaC03)
Carbon
Column
Effluent
105
30
265
40
14
175
347
7.9
27
7.1
2,020
1,150
278
Primary
Cation
Effluent
3.1
0.94
134
20
7.5
185
337
7.5
26
2.7
3,480
344
Primary
An ion
Effluent
3.2
0.98
134
21
7.7
24
167
4.6
18
5.5
830
55
Secondary
Cati on
Ef f 1 uent
0.12
0.01
28
3.9
1.8
26
167
4.7
17
2.9
1,810
193
Secondary
An ion
Ef f 1 uent
0.10
0.01
28
4.0
2.1
1.0
27
1.5
0.49
5.9
163
98
49
Total
Removal
%
100
100
89
90
85
99
92
81
98
92
92
-------
TABLE 11. PERFORMANCE OF THE TWO-STAGE ION EXCHANGE SYSTEM OPERATING ON 1,640 mg/1 TDS FEED WATER
CO
Parameter
(mg/D
Calcium
Magnesuum
Sodium
Potassium
Ammonium
Sul fate
Chloride
Nitrate
Phosphate
PH
(as N)
(as N)
(as PO,,)
Conductivity (ymhos/cm)
TDS
Alkalinity
Acidity
(as CaC03)
(as CaC03)
Carbon
Col umn
Ef f 1 uent
140
47
343
62
33
221
510
2.1
29
7.8
2,740
1,643
409
Primary
Cation
Ef f 1 uent
4.3
1.8
166
35
20
227
495
1.4
28
2.2
4,750
464
Primary
Anion
Effluent
4.3
1.8
166
36
20
18
468
1.7
28
2.6
3,090
232
Secondary
Cation
Effluent
0.24
0.09
63
13
8.3
17
465
1.6
28
2.1
4,760
502
Secondary
Anion
Ef f 1 uent
0.21
0.09
63
14
8.3
1.6
88
0.57
10
6.1
516
238
51
Total
Removal
%
100
100
82
77
75
99
83
72
65
81
86
-------
TABLE 12. EFFECTS OF THE INFLUENT TDS UPON THE TWO-STAGE
ION EXCHANGE SYSTEM PERFORMANCE
610 mg/1 1,150 mg/1 1,640 mg/1
Parameter
TDS TDS TDS
TDS Removal, % 89 92 86
Cation Exchanger
Regeneration Efficiency, % 85 82 85
Anion Exchanger
Regeneration Efficiency, % 90 93 95
Ratio of Brine to Product
Water (Without Brine
Recycling) 0.11 0.22 0.28
Ratio of Brine to Product
Water (With Brine
Recycling) 0.07 0.15 0.18
Estimated Water Recovery*
(With Brine Recycling), % 93 87 85
* Water Recovery (%) = (Product Volume) Y lnn
(Total Feed Volume) A
37
-------
CO
00
100
80
60
40
CO
Q
I-
20
I I
i r
TDS REMOVAL
PRODUCT TDS
1
I I
1
I I
200 400 600
800 1000 1200 1400
INFLUENT TDS, mg/l
500
400 _
300 g
I-
200
100
o
o
a:
a.
1600 1800 2000
Figure 9. Effects of influent TDS upon TDS removal and product TDS.
-------
As indicated in both Table 12 and Figure 9, the two-stage
ion exchange system seemed able to achieve similar degree of IDS
removal and regeneration efficiency under the same typical
operating conditions within the investigated range of influent
IDS values. However, the ratio of the brine volume to the prod-
uct water volume was much higher for 1,640 mg/1 feed water than
the 610 mg/1 feed water. By brine recycling, the brine volume
for each of the three different IDS feed waters could be re-
duced about 30 to 40 percent. This would improve the water re-
covery for the 1,640 mg/1 TDS feed water to the desirable level
of 85 percent.
BRINE CHARACTERISTICS
The ion exchange system generated reject streams from back-
washing, regeneration, and rinsing operations. Among these re-
ject streams, the backwash waste could be reclaimed by recycling
through the pretreatment system (carbon adsorption) to remove
the small amount of suspended solids. However, the reject
streams from the regeneration and rinsing operations contained
undesirable levels of organic materials and inorganic constitu-
ents. These two streams were discharged together and were
classified as the brine from the ion exchange system. An analy-
sis of the brine from a typical operation with the C-20+A-30B+
C-20+A-30B flow scheme on the Pomona carbon-treated secondary
effluent is presented in Table 13. Disposal of the brines from
both cation and anion exchangers was not investigated at Pomona.
Therefore, the cost of brine disposal is not included in the
cost estimates provided in this report although it is recognized
that the cost of brine disposal may be a major component of the
total cost in some particular systems.
Considerable effort was devoted to minimizing the volume of
the various reject streams. As a result of an extensive series
of tests on the Pomona wastewater, which had a normal TDS level
of 610 mg/1, the appropriate volume for each operating sequence
was determined as shown in Table 14. The total reject stream
volume, including backwash, spent regenerant, and rinsing waste,
amounted to about 15 percent of the product water. Since the
backwash water could be recycled, however, disposal of only the
spent regenerant and rinsing waste was necessary. These reject
streams, which were classified as brine, amounted to about 11
percent of the product flow. The brine volumes for other
higher TDS feed waters investigated in this study are shown in
Table 12.
BRINE RECYCLING
There were two major areas of interest in pursuing the
study of brine recycling. The primary interest was to reduce
39
-------
TABLE 13. CHARACTERISTICS OF BRINES FROM THE CATION
AND AN ION EXCHANGE COLUMNS *
Parameter (mg/1)
Cal ci urn
Magnesium
Sodium
Potassium
Ammonium
Sulfate
Chloride
Nitrate
Phosphate
Total Hardness
Alkalinity
Aci di ty
Total Solids
pH
Conductivity
Turbidity
Color
Chemical Oxygen
(as N)
(as N)
(as POO
(as CaC03)
(as CaC03)
(as CaC03)
(ymhos/cm)
(OTU)
(color unit)
Demand (COD)
Cation
Exchange
Columns
500
310
1 ,430
240
280
6,600
30
3
15
2,520
950
9,400
2.2
11,400
2.5
6
15
An ion
Exchange
Col umns
8
0
70
20
2,340
3,200
2,080
100
820
20
1,750
7,590
7.3
17,000
1.5
28
200
* Brine includes spent regenerant, slow rinse water, and fast rinse water.
40
-------
TABLE 14. MINIMUM REJECT STREAMS PER CYCLE OF REGENERATION
Source of Reject Stream
Cation Exchange System
Primary Column
Secondary Column
Combined Flow
Anion Exchange System
Primary Column
Secondary Column
Combined Flow
Sub-Total
Backwash
Water
{gallons)
75
4
4
4
87
Spent
Regenerant
(gallons)
35
5
40
Slow Rinse
Water
(gallons)
75
45
120
Fast Rinse
Water
(gallons)
75
30
105
NOTES:
1. Product water per cycle
2. 1 gallon = 3.8 liters.
= 2,400 gallons (9,120 liters)
-------
the volume of. brine waste to minimize the cost of brine disposal
The secondary interest was to improve the water recovery for the
ion exchange process to reduce the process cost. Some cost
savings for the regenerants could also be secured by a success-
ful brine recycling practice.
During brine recycling study, the regenerant levels for
both the cation and anion exchange columns were maintained at
17.6 g/1 (1.1 Ib/cu ft) of H2S04 and 9.6 g/1 (0.6 Ib/cu ft) of
NH3, respectively.
The typical pilot plant operation flow rates as shown in
Table 3 were used in this study. Extensive studies were first
conducted to characterize the brine during the process of the
regeneration operation. The results of these studies were em-
ployed to decide what portions of the brine flow could be re-
tained for recycling. Figure 10 and Figure 11 represent the
typical distributions of the major impurity ions in the brines
from both cation and anion exchangers, respectively. Generally
speaking, the central portions of both brine streams all con-
tained exceptionally high concentrations of impurity ions.
These were definitely not suitable for recycling. According to
the brine characteristics as illustrated in Figure 10 and Figure
11, a typical zoning diagram such as Figure 12 could be pre-
pared to dictate the various reuses of the recycled brines.
Figure 13 and Figure 14 show the two different modes of re-
generation in terms of the brine handling. The numbers in-
dicated in the parentheses were the average volumes of waters
flowing through those specific lines during each cycle of re-
generation operation. According to these two schematic flow
diagrams, the water recoveries for processes with brine re-
cycling and without brine recycling were about 93 percent and
89 percent, respectively. It is also indicated in the figures
that the total volume of waste brine was greatly reduced from
1000 liters (265 gallons) per cycle to 660 Irters (175 gallons)
per cycle, a reduction of 34 percent, as a result of the re-
cycling of part of the brine produced in the regeneration
operation.
The average qualities of the various effluents from the
two-stage ion exchange system during this study of brine re-
cycling are shown in Table 15. Additional experimental results
are also presented in Table 16 to demonstrate the effects of
brine recycling upon various process parameters. As shown in
both Table 15 and Table 16, the quality of the product water
and the system performance were not deteriorated by brine re-
cycling. The achievement of good product quality and system
performance as well as the improvement in the water recovery
and in the brine waste volume reduction seemed to make the brine
42
-------
4500
4000-
FORMATION
OF CaS04
PRECIPITATE
40
60 80 100 120 140 160
180
ACCUMULATED VOLUME OF BRINE SINCE START OF REGENERATION,
GALLONS (I GALLON = 3.8 LITERS)
Figure 10. Relationship of concentration vs. volume of the impurity
ions in the cation exchanger brine.
43
-------
20,000 r
18,000
16,000
14,000-
v
3»
12,000
510,000
DC
o 8,000
o
o
6,000
4,000
2,000
10 20 30 40 50 60 70 80
ACCUMULATED VOLUME OF BRINE SINCE START OF REGENERATION,
GALLONS (I GALLON= 3.8 LITERS)
Figure II. Relationship of concentration vs. volume of the impurity
ions in the anion exchanger brine.
44
-------
BRINE
OF
ANION
EXCHANGERS
BRINE
OF
CATION
EXCHANGERS
TO BE USED AS RE6ENERANT
W7/W//7A MAKE-UP WATER'
I 1 TO BE DISPOSED WITHOUT RECYCLING
TO BE USED AS PART OF RINSE WATER
0 20 40 60 80 100 120 140 160 180 200
ACCUMULATED VOLUME OF BRINE, GALLONS (I GALLON = 9.8 LITERS)
Figure 12. Zoning diagram of brine recycling.
-------
FEED
WATER
(185)
(TOTAL FLOW)
DEMORALIZATION FLOW
AN-EX REGENERATION FLOW
CAT-EX REGENERATION FLOW
IN GALLONS/CYCLE
(I GALLON = 3.8 LITERS)
WASTE
PRODUCT
Figure 13. Schematic flow diagram for ion exchange process without brine recycling.
-------
FEED
WATER
*- DEMINERALIZATION FLOW
». CAT-EX REGENERATION FLOW
*- ANrEX .REGENERATfON FLOW
(TOTAL FLOW) IN GALLONS/CYCLE
(I GALLON - 3.8 LITERS)
(175) ^
RINSE \ /CATION
WATER/ WASTE
WASTE
•-PRODUCT
i
I
I [SECONDARY
ANION
I
Figure 14. Schematic flow diagram for ion exchange process with brine recycling.
-------
TABLE 15. AVERAGE EFFLUENT QUALITIES OF THE TWO-STAGE ION EXCHANGE
SYSTEM OPERATING WITH BRINE RECYCLING
00
Parameter
(mg/1 )
Cal ci urn
Magnesium
Sodi urn
Potassium
Ammonium (as N)
Sulfate
Chloride
Nitrate (as N)
Phosphate (as P04)
Alkalinity (as CaC03)
Acidity (as CaC03)
TDS
pH
Conductivity (ytnhos/cm)
COD
Carbon
Effluent
59
12
no
13
16
63
116
2.0
29
251
571
7.7
1,000
6.2
Primary
Cation
Effluent
0.41
0.19
49
6.5
7.8
67
114
1.1
29
120
2.9
1,200
Primary
An ion
Effluent
0.43
0.19
49
6.9
8.0
3.6
51
0.58
14
59
6.2
320
Secondary
Cation
Effluent
0.16
0.07
8.1
1.6
1.5
3.3
50
0.60
13
55
3.1
490
Secondary
An ion
Ef f 1 uent
0.15
0.07
8.3
1.5
1.5
0.37
7.3
0.28
0.39
24
74
5.9
64
2.4
Total
Removal
%
100
99
93
89
91
99
94
86
99
87
94
61
-------
TABLE 16. EFFECTS OF BRINE RECYCLING UPON THE TWO-STAGE
ION EXCHANGE PILOT PLANT PERFORMANCE
Parameter
Regeneration
with Brine
Recycling
Regeneration
without Brine
Recycling
TDS Removal, %
87
89
Water Recovery, %
93
89
Regeneration Efficiencies,
Cation Exchanger
Anion Exchanger
87
85
90
Total Ion Leakages, %
Cation Exchanger
Anion Exchanger
5.2
4.8
5.8
6.5
49
-------
recycling a very attractive and very beneficial practice. How-
ever, the actual savings in the overall process cost (excluding
brine disposal) might not be too significant because of some
additional expenses required by the additional storage capacity
for holding the brine, and the higher degree of automation for
recycling operation.
Table 17 shows the average characteristics of the different
portions of the brine in various reuses. The determinations of
the TDS values for samples E and F, which had very high concen-
trations of ammonium and chloride ions, were found to be quite
inaccurate using the routine analytical procedures. Therefore,
the calculated values instead of measured values are shown in
Table 17.
RESIN LIFE
Since the resin life is a major cost-related parameter in
the ion exchange process, a great deal of research effort was
expended in the area of resin life and resin stability.
During the entire period of experimentation with the
Infilco two-stage ion exchange pilot plant, resin samples were
withdrawn at various intervals for inspection and analysis. The
resin analysis was conducted by the Infilco analytical labora-
tory at Tucson, Arizona. The last set of resin samples analyzed
by Infilco were taken after a total of 450 operation runs,
approximately 30 months of operation. The total volume of water
treated by the resins at the time of last sampling was about
4.73 million liters (1.25 million gallons). Table 18 shows the
results of the dry screen analysis of the cation and anion .ex-
change resins. Table 19 shows the resin capacity of the cation
and anion exchange resin in terms of percent of the capacity of
virgin resins operating under identical experimental conditions.
The rinse water requirements for the anion exchange resins in
various operation stages were summarized in Table 20.
As indicated in Table 19 and Table 20, the primary anion
exchange resin, Duolite A-30B, showed substantial deterioration
in both resin capacity and rinse efficiency. The secondary
anion exchange resin, Duolite A-30B, also showed a certain de-
gree of degradation. These results as determined by standard
laboratory procedures seemed to be very discouraging. However,
the laboratory procedures employed regeneration levels many
times higher than that employed in the actual pilot plant
operation. In the pilot plant, the anion exchange resins were
regenerated with ammonium hydroxide solution at the very low
level of 9.6 g/1 (0.6 Ib/cu ft) of NH3, while in the laboratory
resin analysis, the anion exchange resins were regenerated with
caustic soda at an excessively high regeneration level.
50
-------
TABLE 17. QUALITY CHARACTERISTICS OF THE VARIOUS PORTIONS OF THE BRINES
Parameter
mg/1
Calcium
Magnesium
Sodi urn
Potassium
Ammonium (as N)
01 Sulfate
i— »
Chloride
Nitrate (as N)
Phosphate (as POJ
Alkalinity (as CaC03)
Acidity (as CaC03)
TDS
pH
Conductivity (pmhos/cm)
COD
(A)*
4% H2S04
Make-up
Water
6.7
2.4
127
13
17
247
87 -•'
0.80
29
38
525
3.2
1,100
17
(B)*
Cation
Exchanger
Rinse
Water
59
12
375
46
47
1,960
50
0.35
13
899
2,920
2.1
7,110
4.5
(C)*
Cation
Exchanger
Waste
Water
1,210
286
2,070
224
289
10,800
81
1.0
55
2,510
16,000
1.9
26,900
16
(D)*
4% NHitOH
Make-up
Water
0.38
0.24
57
5.6
16
9.4
98
0.68
49
2
222
4.3
514
12
(E)*
An ion
Exchanger
Rinse
Water
0.55
0.77
117
13
395
254
627
8.3
128
356
1 ,380**
7.7
3,720
21
(F)*
An ion
Exchanger
Waste
Water
0.19
0.08
93
7.4
4,498
4,370
5,350
63
1,480
1,800
12,600**
7.4
28,200
162
* Taken from the same streams as indicated in Figure 14. ** Calculated values.
-------
TABLE 18. DRY SCREEN ANALYSIS OF THE CATION AND ANION EXCHANGE RESINS (% RETAINED CUMULATIVE)
en
Analysis
U. S. Screen No.
16
20
30
40
50
70
Pan
Bead Rating (%)
Perfect
Crazed
Broken
Pitted
Cation
Virgin
Resin
1 - 6
21 - 43
67 - 84
82 - 98
97 - 99
100
-
99 - 100
—
—
• _ ~
Exchange Resin, C - 20
Primary
Cation
Col umn
Resin*
0.9
22.5
48.3
89.6
98.8
100
Trace
99 - 100
_ —
Trace
«H_ «-
Secondary
Cation
Column
Resin*
1.4
13.3
49.4
90.0
98.5
100
Trace
99 - 100
—
Trace
w «• —
Anion Exchange Resin,
Virgin
Res i n
—
0.6
16.9
54.4
87.5
100
-
99-100
—
—
Primary
Anion
Col umn
Resin*
—
7.4
33.8
72.2
93.2
100
Trace
99 - 100
—
Trace
A - 30 B
Secondary
Cation
Col umn
Resin*
—
7.0
31.3
68.7
90.7
100
Trace
99 - 100
—
Trace
* Resin samples were taken after 450 demineralization cycles
-------
TABLE 19. RESIN CAPACITIES OF THE CATION AND
ANION EXCHANGE RESINS
Resin Sample Resin Capacity
Primary Cation Column Resin 100%
Secondary Cation Column Resin 100%
Primary Anion Column Resin 62.3%
Secondary Anion Column Resin 75.6%
NOTES:
1. Resin capacity is expressed as the % of the capacity of the
virgin resin operating under identical conditions in
laboratory test.
2. Capacity based on water produced with a conductivity equal
to or better than 100 ymhos/cm.
3. Resin samples were taken after 450 demineralization cycles.
53
-------
TABLE 20. RINSE WATER REQUIREMENTS OF ANION RESINS
Resin Sample
100 ymhos/cm
Effluent
(gal/cu ft)
10 ymhos/cm
Effluent
(gal/cu ft)
Virgin Resin
50
Primary Anion Column Resin
45
213
Secondary Anion Column Resin
15
50
NOTES:
1. Duolite A-30B intermediate base anion exchange resin was
used in this study.
2. The resin samples were taken after 450 demineralization
cycles.
3. 1 gal/cu ft = 135 1/cu m.
54
-------
Therefore, the laboratory tests could not be directly translated
to actual operations. As far as the pilot plant operation was
concerned, there was no discernible decrease in performance
efficiency throughout the entire pilot plant operation period.
The best indication of the resin life probably lies in the
pilot plant performance itself. Consequently, a 15 1pm (4 gpm)
automated single-stage ion exchange pilot plant was designed and
operated to handle the special study on resin life. Only a
single-stage instead of a complete two-stage system, as in the
Infilco two-stage pilot plant, was employed in this pilot plant
to reduce the cost for the complex automation. Figure 15 shows
the general layout of the automated single-stage system. This
system contained two identical reinforced fiberglass resin
columns. Each column had a 30.5 cm (12 in) diameter and was
122 cm (48 in) in height. The cation exchange column, which con-
tained 56.6 liters (2 cu ft) of Duolite C-20 resin, preceded the
anion exchange column which had 56.6 liters (2 cu ft) of Duolite
A-30B resin. The system was designed to provide four regenera-
tions of the cation exchange column to one regeneration of the
anion exchange column over a repeating time cycle of 720 min-
utes. Each 180 minutes of the 720 minute time cycle the cation
exchange column would regenerate and each 720 minutes of the
720 minute time cycle, the anion exchange column would regen-
erate. This repeating time cycle was controlled by a programmed
cam timer, which automatically made one complete revolution
every 720 minutes.
The regeneration levels for the cation and anion exchange
resins were maintained at the same values of 17.6 g/1 (1.1 Ib/cu
ft) of H2S04 and 9.6 g/1 (0.6 Ib/cu ft) of NH?, respectively.
The system operated under these selected conditions produced an
average TDS reduction of 70 percent over the entire operation
period of 32 months. The actual monthly fluctuation of the TDS
removal is shown in Figure 16. Some operation statistics are
shown in Table 21 to compare this automated single-stage system
with the Infilco two-stage system. A typical analysis of the
feed (carbon-treated secondary effluent) and product water
qualities is shown in Table 22. As indicated in Table 21, the
automated single-stage ion exchange pilot plant was operated
under much more intensive regeneration conditions, and had
treated much more wastewater than the Infilco two-stage ion ex-
change pilot plant. However, the system performance, as re-
flected by the fairly consistent TDS removal throughout the en-
tire operation period, did not show any significant deteriora-
tion. Although the question of long-term effects of wastewater
upon the resins has not been completely answered, an optimistic
outlook appears justified. Liberal allowances are included in
the cost estimates*for resin replacement, 10 percent per year
for the cation exchange resin and 20 percent per year for anion
55
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en
4% H2S04
TANK
4% NH4OH
TANK
INFLUENT WATER
(FROM CARBON COLUMN)
CARBON
WATER
SURGE
TANK
SLOW-RINSE LINE
._ JZ IZZI — —•=L"T "1
AuunwiA SLOW-RINSEl '
AMMONIA i IMC '
LINE
\
\
BACKWASH a FAST-RINSE
PUMP
ACID
PUMP
BACKWASH a FAST-RINSE
PUMP
FLOW
INDICATOR
CATION
EXCHAN
ANION
XCHAN6E
CONDUCTIVITY
PROBE
FLOW
METER
INFLUENT
PUMP
OVER-FLOW
TO WASTE
PRODUCT WATER
STORAGE TANK
TO WASTE
Figure 15. General layout of the automated single-stage ion exchange pilot plant.
-------
en
100
90
80
70
60
50
UJ
s 40
30
20
10
/—t
AVERAGE
I
1
10 15 20
MONTHS ON STREAM
25
30
Figure 16. System performance of the single-stage ion exchange unit.
35
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TABLE 21. OPERATION STATISTICS OF SINGLE-STAGE VS.
TWO-STAGE ION EXCHANGE SYSTEM
Parameter
Automated
Single-Stage
System
Infilco
Two-Stage
System
Total Operation Period, months
32
48
Total Volume of Water Treated
Cation Exchange System
million gallons/cu ft
million liters/cu m
Am" on Exchange System
million gallons/cu ft
million liters/cu m
2.3
311
2.3
311
0.2
27
0.6
81
Total Regeneration Performed, cycles
Cation Exchange System
Anion Exchange System
7,616
1,904
744
744
58
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TABLE 22. TYPICAL PERFORMANCE OF THE AUTOMATED
SINGLE-STAGE ION EXCHANGE SYSTEM
Parameter
Calcium
Magnesium
Sodi urn
Potassi urn
Ammonium (as N)
Sul fate
Chloride
Nitrate (as N)
Phosphate (as POi*)
pH
Conductivity (ymhos/cm)
TDS
Alkalinity (as CaC03)
Feed
Water
(mg/1)
48
8.2
76
11
12
63
70
0.93
27
7.5
910
439
233
Product
Water
(mg/D
0.50
0.10
27
4.6
6.9
16
27
0.40
5.9
6.5
236
127
61
Total
Removal
%
99
99
64
58
43
75
61
57
78
74
71
59
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exchange resin, but resin replacement at these rates will still
amount to less than 10 percent of the total cost of the ion ex-
change process.
STEADY-STATE PILOT PLANT OPERATION
A series of 80 complete operating cycles was conducted on a
routine operating basis on the Infilco two-stage ion exchange
unit to demonstrate a steady-state operation. The unit was
operated on what were considered optimum regeneration conditions
as determined from the various aforementioned special investi-
gations. The various operating flow rates were restricted by
the physical conditions of the pilot plant, and were regulated
to the typical pilot plant operating flow rates as shown in
Table 3. The TDS values of the feed and product waters for each
operating cycle are shown in Figure 17. The overall average TDS
values for feed and product waters are 610 mg/1 and 72 mg/1, re-
spectively, for a reduction of about 89 percent.
The average water quality of each of the various effluents
from the ion exchange pilot plant is shown in Table 23. As in-
dicated in Table 23, the primary cation exchange column removed
almost completely the calcium and magnesium ions. However, only
about half of the concentrations of the sodium, potassiurn, and
ammonium ions were removed in this column. The remaining half
of these less selective monovalent cations were effectively re-
moved by the secondary cation exchange column, which had a much
higher state of regeneration as a result of the semi-counter-
current regeneration flow pattern as indicated in Figure 3.
Similarly, the primary anion exchange column removed most of the
sulfate ions while the nitrate, chloride, and phosphate ions
were partially removed by both the primary and secondary
columns .
The product water pH of 5.8 was lower than desirable but
it was anticipated that the product of the ion exchange process
would be blended with non-demineralized wastewater prior to re-
use. The buffer capacity of the ion exchange product water was
quite low when compared to the water with which it would be
blended, and some laboratory tests confirmed that the pH of the
ion exchange product water would increase to the pH of the water
with which it was blended. Otherwise, an aerating tower can be
employed to strip out the carbon dioxide from the ion exchange
product to raise the pH to a nearly neutral value.
During the steady-state pilot plant operation period, the
entire system functioned satisfactorily with no operational /
problems. This has demonstrated ion exchange as a reliable and
practical process for wastewater demineralization.
60
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CT>
700
^600
o>
E
•»
V)
o50O
o
CO
4
O
UJ
400
o
CO
CO
300
200
100
PRODUCT
10
20
70
80
30 40 50 60
OPERATING CYCLE
Figure 17. Performance of the two-stage ion exchange system under optimum
operating conditions.
90
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TABLE 23. AVERAGE WATER QUALITY CHARACTERISTICS OF THE TWO-STAGE
ION EXCHANGE PILOT PLANT UNDER OPTIMUM OPERATING CONDITIONS
Parameter
(mg/1 )
Calcium
Magnesium
Sodium
Potassium
Ammonium (as N)
Sulfate
Nitrate (as N)
Chloride
Phosphate (as POJ
Alkalinity (as CaC03)
Acidity (as CaC03)
PH
Total COD
Silica (as Si02)
Conductivity (pmhos/cm)
TDS
Secondary
Carbon Primary Primary Secondary Anion
Column Cation Anion Cation Column
Effluent Column Column Column Effluent
(Feed) Effluent Effluent Effluent (Product)
53 2.0
17 0.59
126 61
14 7.3
20 9.6
72 72
2.9 2.8
135 132
27 27
218
no
7.4 2.7
10 9.7
23
1,040 1,390
610 298
1.7
0.56
59
7.1
9.2
3.6
1.6
84
15
51
5.7
6.8
390
198
1.1
0.38
16
1.9
4.0
3.6
1.5
83
14
110
2.8
5.5
1,040
104
0.60
0.0
15
1.9
3.8
1.3
0.35
14
0.25
39
5.8
3.7
23
100
72
62
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SECTION 6
COST ESTIMATE
Cost estimates have been prepared for the ion exchange pro-
cess based upon the Infilco two-stage ion exchange pilot plant
study conducted at Pomona Research Facility. Figure 18 shows
the effect of plant size on the process cost for a 90 percent
TDS removal from 600 mg/1 feed to 60 mg/1 product. This cost
curve clearly demonstrates that the process cost approaches its
minimum level when the plant size is equal to or larger than
37,850 cu m/day (10 MGD). Therefore, a 37,850 cu m/day (10 MGD)
plant is used as a standard size for various cost estimates pre-
sented in this report. A detailed analysis on the process cost
estimate for a 37,850 cu m/day (10 MGD) product water ion ex-
change plant is shown in Table 24. The various assumptions em-
ployed in making this cost analysis are also included in the
table. This cost estimate corresponds to the costs as of
August, 1973.
As indicated in Table 24, the total process cost of
5.80/1,000 liters (22.40/1,000 gallons) is comprised of 0.90/
1,000 liters (3.70/1,000 gallons) for capital amortization and
4.90/1,000 liters (18.70/1,000 gallons) for operation and main-
tenance. The dominant cost item is the regeneration expenses,
with 3.5^/1,000 liters (13.60/1,000 gallons) attributed to this
item. The 5.80/1,000 liters (22.40/1,000 gallons) cost esti-
mate covers only the cost of the ion exchange process itself but
not the cost of pretreatment (it will cost an additional 2.30/
1,000 liters or 90/1,000 gallons for carbon adsorption) and
brine disposal.
Since most demineralization applications will involve feed
water of higher TDS than that encountered at Pomona, some cost
estimates have been made to cover an influent TDS range as high
as 1,500 mg/1, which is considered to be the practical cost-
effective limit for ion exchange applications to wastewater de-
mineralization. To make such estimates it is necessary to
assume the influent concentrations of the various ionic con-
stituents in the higher TDS waters. The cost estimate could be
different if the individual chemical constituents would vary
from the assumed concentrations. However, the use of TDS as a
gross parameter is sufficient to obtain reasonably close esti-
mates. Figure 19 shows the effect of the influent TDS on the
63
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INFLUENT IDS * 600 mg/l
1 M6D - 3,785 cu m/day
1^/1000 GALLONS
= 0.26*/IOOO LITERS
2 5 10
PLANT SIZE, MGD
20
50
100
Figure 18. Effect of plant size on id? exchange process cost- with 90%TDS reduction.
-------
TABLE 24. COST ESTIMATE OF TWO-STATE ION EXCHANGE PROCESS
37,850 cu m/day (10 MGD) PRODUCT WATER PLANT
Amortization of Capital <£/l,000 Gallons
$1,660,000; 20 years @ 5% 3.7
Operation and Maintenance
Regeneration
Cation Exchanger (Sulfuric Acid) 10.0
Anion Exchanger (Ammonium Hydroxide) 3.6
Resin Replacement 2.0
Maintenance Materials 1.0
Power 0.8
Labor 1.3
18.7
Total Process Cost 22.4
Assumptions:
1. Influent TDS = 600 mg/1; effluent TDS = 60 mg/1
2. Water recovery = 89%
3. Regeneration efficiency : cation = 85%; anion = 90%
4. Annual resin replacement : cation = 10%; anion = 20%
5. U/1,000 gallons = 0.26^/1,000 liters
6. Based on August, 1973 material and construction costs
7. Costs for brine disposal and pretreatment are not included
65
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CO
z
o
_l
_J
CD
O
O
O
o
o
CO
CO
LJ
o
o
a:
a.
I
100
90
80
70
60
50
40
30
20
10
I MGD = 3785 cu m/day
I*/IOOO GALLONS = 0.26*71000 LITERS
10 MGD
500 1000 1500 2000
INFLUENT TDS,mg/l
Figure 19. Effect of influent TDS upon ion exchange
process cost—with 90% TDS reduction.
66
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total process cost for plant sizes of 3,785 cu m/day (T MGD)
and 37,850 cu m/day (10 MGD), with 90 percent IDS reduction.
Additional cost estimates have been made to compare the
two-stage ion exchange process with the single-stage ion ex-
change process, which was operated at the Pomona Research
Facility solely for resin life study purposes. The single-
stage system was operated under the same optimum regeneration
conditions as the two-stage system to simulate the full scale
operation of the primary stage of the two-stage system. Con-
sequently, the single-stage system achieved only about 70 per-
cent IDS removal which was close to the performance of the pri-
mary stage of the two-stage system, whereas the complete two-
stage system achieved about 90 percent IDS removal. Table 25
shows the total process costs for both single-stage and two-
stage systems to produce product waters with approximately 70
percent and 90 percent TDS reductions, respectively. As in-
dicated in this table, the total process cost of 4.6<£/l,000
liters (17.8^/1,000 gallons) for the single-stage system is
about 20 percent lower than the 5.8(^/1,000 liters (22.4^/1,000
gallons) for the two-stage system. However, if the product
waters from these systems are to be blended with non-deminer-
alized water to achieve one-third TDS reduction from 600 mg/1
in the influent waters to 400 mg/1 in the blended waters, then
the unit process cost for the two-stage systems will be
slightly less than the single-stage system as shown in Table 26.
Table 27 illustrates another case of cost estimate compari-
son between the single-stage system and the two-stage system.
In this case, the TDS of the influent is 1,000 mg/1 and the
TDS of the blended water set at 500 mg/1. Contrary to pre-
vious case as shown in Table 26, the single-stage system in-
stead of the two-stage system is the cheaper process to accom-
plish this particular objective.
Although there are slight differences in the cost esti-
mates as indicated in Table 26 and Table 27 for a different
system design, it seems the differences are so little that both
single-stage and two-stage systems can be considered equally
feasible for wastewater demineralization whenever a blending
operation can be used in the system to obtain a blended product
water to meet a less restrictive TDS removal requirement. How-
ever, the two-stage ion exchange process is more economical
than the single-stage process in producing an unblended product
water with a low effluent TDS.
67
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TABLE 25. COST ESTIMATE COMPARISON OF TWO-STAGE VS. SINGLE-STAGE
ION EXCHANGE SYSTEM - 37,850 cu m/day (10 MGD) PLANT
Single - Stage Two-Stage
(fc/1.000 Gallons ^/l .000 Gallons
Amortization of Capital
20 years @ 5% 3.3 3.7
Operation and Maintenance
Regeneration
Cation Exchanger (H^OJ 7.7 10.0
Anion Exchanger (NH4OH) 2.3 3.6
Resin Replacement 1.9 2.0
Maintenance Materials 0.9 1.0
Power 0.7 0.8
Labor _ 1 .0 _ 1 -3
Total Process Cost 17.8 22.3
Assumptions:
1. TDS of feed water = 600 mg/1
2. TDS of product water : single-stage = 200 mg/1; two-stage = 60 mg/1
3. Water recovery = 89%
4. Regeneration efficiency : cation = 85%; an ion = 90%
5. Resin life : cation = 10 years; an ion = 5 years
6. Costs for pretreatment and brine disposal are not included
7. ltf/1,000 gallons = 0.26
-------
TABLE 26. COST ESTIMATE COMPARION OF TWO-STAGE
VS. SINGLE-STAGE ION EXCHANGE SYSTEM
WITH BLENDING OPERATION
.000 gal Ions
37,850 cu m/day (10 MGD) Blended Product
Water Flow - with Different Ion
Exchange Plant Size
a. Single-stage System - 18,930 cu m/day
(5 MGD) Plant 9.7
b. Two-Stage System - 14,000 cu m/day
(3.7 MGD} Plant 9.4
B. 37,850 cu m/day (10 MGD) Ion Exchange
Plant - with Different Blended Product
Water Flow
a. Single-Stage System - 75,700 cu m/day
(20 MGD) Blended Product Water Flow 9.1
b. Two-Stage System - 102,000 cu m/day
(27 MGD) Blended Product Water Flow 8.5
Assumptions:
1. TDS of feed water = 600 mg/1
2. TDS of blended product water = 400 mg/1
3. Regeneration efficiency : cation = 85%; anion = 90%
4. Resin life : cation = 10 years; anion = 5 years
5. U/1,000 gallons = 0.26
-------
TABLE 27. COST ESTIMATE FOR PRODUCING 500 mg/1 TDS
BLENDED PRODUCT WATER - 37,850 cu m/day (10 MGD)
ION EXCHANGE PLANT SIZE
n . Single-stage
Parameter System
Total Blended Product Flow
MGD 13.5
cu m/day 51 ,100
Total Process Cost
i/1 ,000 gallons 19.8
-------
REFERENCES
1.
2.
3.
Kunin, Robert, and Donald G. Downing, "New Ion Exchange
Systems for Treating Municipal, Domestic, and Industrial
Waste Effluents," paper presented at the International
Water Conference, Pittsburgh, Pennsylvania, October,
1970.
Technical Bulletin (April, 1972), ICI Australia Limited,
Melbourne, Australia.
Duolite Tech Sheet No. 120, (October, 1968) and Duolite
Data Leaflet No. 35 (June, 1969), Resinous Products
Division, Diamond Shamrock Chemical Co., Redwood City,
CA 94063.
4.
Technical Bulletin No. IE-119-67 (August, 1967)
and Haas, Philadelphia, Pennsylvania 19105.
Rohm
71
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
EPA-600/2-77-146
2.
3. RECIPIENT'S ACCESSION-NO.
riTLE AND SUBTITLE
Wastewater Demineralization by Two-Stage Fixed-Bed
Ion Exchange Process
5. REPORT DATE
September 1977(Issuing Date)
6. PERFORMING ORGANIZATION CODE
AUTHOR(S)
Ching-lin Chen
Robert P. Miele
8. PERFORMING ORGANIZATION REPORT NO
'ERFORMING ORGANIZATION NAME AND ADDRESS
County Sanitation Districts of Los Angeles County
Whittier, California 90607
10. PROGRAM ELEMENT NO.
1BC611
11. CONTRACT/GRANT NO.
14-12-150
12. SPONSORING AGENCY NAME AND ADDRESS
Municipal Environmental Research Laboratory—Cin., OH
Office of Research and Development
U.S. Environmental Protection Agency
Cincinnati. Ohio 45268
13. TYPE OF REPORT AND PERIOD COVERED
Final
14. SPONSORING AGENCY CODE
EPA/600/14
15. SUPPLEMENTARY NOTES
Project Officer: Irwin J. Kugelman (513-684-7631)
16. ABSTRACT
A 9.5 1/min (2.5 gpm) two-stage fixed bed ion exchange process (primary cation -
primary anion - secondary cation - secondary anion) was operated on a feed of
carbon treated secondary effluent for 48 months at Pomona, California. To achieve
high levels of regeneration efficiency regenerant was passed counter current to the
feed, and regenerant levels were held to 17.6 g HoSOA per liter of cation resin
(1.1 lb/ft3) and 9.6 g NH3 per liter of anion resin (0.6 lb/ft3). At this level
regenerant efficiency was 85% for the cation resin and 90% for the anion resin.
TDS removal for feed range of 600 mg/1 to 1700 mg/1 was in excess of 90% despite
the low level of regenerant used.
A single stage system (primary cation - primary anion) was set us to determine resin
lifetime performance. The same feed and regenerant was used as in the two-stage
system. During a 32 month period the cation resin was regenerated over 7,000
times and the anion resin was regenerated almost 2,000 times. No evidence of any
deterioration was observed. The system consistently achieved 70% TDS removal.
Estimates are presented on treatment costs as a function of plant capacity and
feed TDS based on the data generated.
7.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Demineralizing
Desalting
Ion exchanging
Purification
Water reclamation
Wastewater Renovation
13B
8. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (ThisReport/
Unclassified
21. NO. OF PAGES
82
20. SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
72
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