EPA-600/2-77-152
September 1977
WASTEWATER DEMINERALIZATION BY CONTINUOUS
COUNTER-CURRENT 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 Re-
search Laboratory, U.S. Environmental Protection Agency, and ap-
proved 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.
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FOREWORD
The Environmental Protection Agency was created because of in-
creasing public and government concern about the dangers of pollu-
tion to the health and welfare of the American people. Noxious
air, foul water, and spoiled land are tragic testimony to the de-
terioration of our natural environment. The complexity of that
environment and the interplay between its components require a
concentrated and integrated attack on the problem.
Research and development is that necessary first step in problem
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 wastewater and solid
and hazardous waste pollutant discharges from municipal and com-
munity sources, for the preservation and treatment of public
drinking water supplies, and to minimize the adverse economic,
social, health, and aesthetic effects of pollution. This pub-
lication is one of the products of that research; a most vital
communications link betweeen the researcher and the user community.
Renovation of wastewater for recycle and reuse may require partial
demineralization of effluents from convention treatment. This
report summarizes studies of demineralization of secondary effluent
by a unique ion exchange process. The technique employed provides
continuous counter-current contact between the wastewater and the
ion exchange resins which provides potential economies over the
conventional contacting techniques.
Francis T. Mayo, Director
Municipal Environmental Research
Laboratory
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ABSTRACT
A wastewater demineralization study employing a 38 1pm
(10 gpm) continuous counter-current ion exchange pilot plant,
manufactured by the Chemical Separations Corporation, Oak Ridge,
Tennessee, has been conducted at the County Sanitation Districts'
Pomona Research Facility, Pomona, California. The study was
jointly funded by the U.S. Environmental Protection Agency and
the County Sanitation Districts of Los Angeles County,
The continuous counter-current ion exchange pilot plant
has demonstrated a promising regeneration efficiency for both
cation and anion exchangers. The brine volume produced by the
process was approximately 8 percent of the product flow, thus
yielding a 92 percent water recovery. The annual resin operation
losses were about 5 percent for the cation exchanger and 15 per-
cent for the anion exchanger. These high resin losses, however,
account for less than 5 percent of the total process cost.
A cost estimate for a 37,850 cu m/day (10 MGD) continuous
counter-current ion exchange plant based on Pomona pilot plant
operating results has been made. The estimated total process
cost of 4.8^/1,000 liters (18.3^/1,000 gallons) was based on the
use of carbon-treated secondary effluent with an average TDS
concentration of 600 mg/1 to produce a product water with 82 per-
cent reduction in TDS.
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 Pro-
tection Agency. Work on this report was comoleted as of August
1972.
IV
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CONTENTS
Foreword iii
Abstract iv
Figures vi
Tables vii
Acknowledgment viii
1. Introduction 1
2. Conclusions 3
3. Recommendations 4
4. Process and Pilot Plant Descriptions 5
5. Pilot Plant Operation 12
6. Results and Discussions 20
Removal of ionic impurities 20
Regeneration Efficiencies 22
Brine characteristics 25
Resin stability 26
Process reliability 30
7. Process Cost Estimate 32
References 35
Appendix A 36
Appendix B 38
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FIGURES
Number Page
1 Photograph of Chem-Seps continuous counter-
current ion exchange pilot plant 7
2 Schematic diagram of Chem-Seps continuous
counter-current ion exchange contactor 8
3 General layout of the Chem-Seps continuous
counter-current ion exchange pilot plant 13
4 Layout of pilot plant sample points 15
5 Performance of the Chem-Seps continuous
counter-current ion exchange system 31
VI
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TABLES
Number Page
1 Resin Properties and Characteristics 11
2 Typical Pilot Plant Performance Using
H2S01+-Ca(OH)2 Regeneration Mode 21
3 Pilot Plant Best Performance Using
H2SOi4-NHitOH Regeneration Mode 23
4 Average Water Characteristics During Steady-State
Pilot Plant Operation (H2SO,-NH^OH Regeneration
Mode) 24
5 Average Brine Characteristics During Steady-State
Pilot Plant Operation (H2S0lt-NHlfOH
Regeneration Mode) 27
6 Typical Screen Analysis of Resin Particle
Size Distribution 28
7 Particle Size Distributions of the Resins
from Resin Recovery Tanks 29
8 Process Cost Estimate 33
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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, I. R, Higgins of
Chemical Separation Corporation, Oak Ridge, Tennessee, for his
advice and cooperation in this effort.
Great appreciation is also extended to the operating and
laboratory staff of the Pomona Advanced Wastewater Treatment
Research Facility.
vm
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SECTION 1
INTRODUCTION
An extensive wastewater demineralization study has been
conducted since 1967 at the County Sanitation Districts' Pomona
Research Facility, which is jointly funded by the U.S.
Environmental Protection Agency and the County Sanitation
Districts of Los Angeles County. The study has covered the re-
verse osmosis, electrodialysis and ion exchange processes. The
preliminary results of this extensive study indicated that the
ion exchange process was most promising for wastewater demin-
eralization based on both economical and technical considera-
tions(l). The ion exchange pilot plant employed in previous
wastewater demineralization studies was a conventional fixed-bed
system(2). Recently some moving-bed ion exchange systems which
have been shown to be more efficient than the fixed-bed system
in both municipal and industrial water treatments are available
on the market. Therefore, the wastewater demineralization study
has been expanded to include the moving-bed ion exchange process
The results of this special study are presented in this report.
A continuous (moving-bed) counter-current ion exchange pro-
cess has the following inherent advantages and disadvantages as
compared with the conventional fixed-bed ion exchange process.
A. Advantages:
1. A smaller resin inventory is required.
2. Resin regeneraton is more efficient.
3. Ion leakage is lower at a given level of regenerant,
4. Brine disposal problem is minimized.
B. Di sadvantages :
1. More sensitive and accurate automation system for
efficient process control is required.
2. Higher operating pressure is required.
3. Resin attrition loss is greater.
1
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This study has been conducted to determine if the advan-
tages outweigh the disadvantages for the demineralization of
wastewater. More specifically, the objectives of this study are
(a) to optimize the operating parameters of a continuous counter-
current ion exchange system for wastewater demineralization; and
(b) to develop the process cost for wastewater demineralization
by the continuous counter current ion exchange system.
Because of the dynamic nature of the system operation, it
requires a great deal of balancing and adjustment in a continu-
ous counter-current ion exchange process to achieve an optimum
performance. Therefore, a high degree of flexibility and accur-
acy in the various functional controls is very essential for a
successful process demonstration. Unfortunately, the continu-
ous counter-current ion exchange pilot plant provided by the
Chemical Separations Corporation (Chem-Seps) of Oak Ridge,
Tennessee, was not well equipped with such necessary features
in its original design. Consequently, a series of major cor-
rections and modifications had been done on the pilot plant be-
fore a six month period of intensive process evaluation study
was satisfactorily initiated in March, 1972.
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SECTION 2
CONCLUSIONS
The following conclusions can be drawn from this plot
plant study:
A. The continuous counter-current ion exchange process
requires a more sophisticated process control system than the
conventional fixed-bed co-current ion exchange process. Par-
ticularly, an accurate control of the resin pulsing operation
is very essential for the success of the process.
B. The continuous counter-current ion exchange process
achieved an 88 percent regeneration efficiency for the cation
exchanger and 90 percent efficiency for the anion exchangers
which were slightly better than those obtained for a two-stage
fixed-bed ion exchange process operating on the same wastewater
with same levels of regenerant dosages.
on
C. The monovalent ion removal achieved by the continuous
iun exchange system in this study was poorer than the fixed-bed
ion exchange system in the previous study.
D. The brine volume produced by the continuous counter-
current ion exchange process was about 8 percent of the product
flow, thus yielding a 92 percent water recovery.
E. The actual resin losses caused by the plant operation
were estimated to be about 1.5 percent for the cation exchange
resin and 4.5 percent for the anion exchange resin during the
2,760 hours of on-stream operation.
F. The cost for the process to achieve a 82 percent TDS
reduction from a 600 mg/1 TDS feed water in a 37,850 cum/day (10
MGD) plant is estimated to be 4.8^/1,000 liters (18.3^/1,000
gallons) of product water. The costs for the carbon adsorption
pretreatment and the brine disposal are not included in the pro-
cess cost estimate.
R. The total process cost, including the carbon adsorp-
tion pretreatment, for a blended product water with 500 mg/1 TDS
is about 3.3
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SECTION 3
RECOMMENDATIONS
Although some shortcomings pertaining to the mechanical
design of the Chem-Seps pilot plant have prevented the studies
from being completely successful, these studies have revealed
some promising results. These results have well justified pur-
suing a research effort on the continuous counter-current ion
exchange process for wastewater demineralization.
The following objectives are proposed to be achieved by
the future pilot plant studies:
A. To fully demonstrate the capability of the continuous
counter-current ion exchange process in controlling the leakage
of the monovalent ions;
B. To devise some effective means of rinse control to
prevent the undesirable contamination of the product water by
the excessive regenerants;
C. To overcome the problems associated with the lime
slurry application;
D. To demonstrate the long-term reliability of the pro-
cess performance;
E. To study the effects of the high TDS feed water upon
the process performance and process economy;
F. To study the attrition effects of the process opera-
tion upon the resin life on a long-term operation basis; and
G. To study the feasibility of recovering the ammonium
ion from the brine, in the form of the valuable ammonium ni-
trate, by adopting a new HN03-NH40H regeneration scheme.
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SECTION 4
PROCESS AND PILOT PLANT DESCRIPTIONS
PROCESS FEATURES
The basic ion exchange reactions involved in an ion ex-
change process are practically the same for the fixed-bed
co-current and the continuous counter-current ion exchange
systems. However, the reaction rate and process efficiency
are somewhat different between the two systems. The continuous
counter-current ion exchange system may be more efficient due
to the special features employed in its system design. These
special features and their functions are described as follows:
A. The system can simultaneously perform the four basic
operating cycles - demineralization, backwash, regeneration and
rinse - in a single loop column separated by control valves.
This multiple operation function greatly reduces the necessary
resin inventory and thus the capital cost.
B. The flow of the regenerated resin is opposite to the
flow of the feed water, which is in a downflow pattern. This
counter-current flow feature insures that the last resin the
feed water being contacted with is in the freshest stage of re-
generation, thus making effective the removal of the monovalent
ions which are at the lower end of the selectivity list, such
as sodium, potassium, ammonium, chloride and nitrate ions.
C. The regenerant flow is counter to the flow of the ex-
hausted resin. The exhausted resin is introduced into the re-
generation section from the outlet end of the regenerant waste.
Therefore, the resin can be progressively regenerated by
stronger regenerant as it flows further through the regeneration
section. This flow pattern can achieve the following objec-
ti ves :
1. Minimizing the formation of calcium sulfate pre-
cipitates in the resin bed when sulfuric acid is employed as re-
generant for cation exchanger.
2. Minimizing the handling problem when the low cost
lime slurry is used as the regenerant for anion exchanger.
3. Maximizing the utilization of regenerant.
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D. The continuous flow function of the process allows the
brine stream to be more uniform in terms of quantity and quali-
ty, thus minimizing the brine disposal problems.
PILOT PLANT DESCRIPTION
General
The Chem-Seps continuous counter-current ion exchange pi-
lot plant contained all the aforementioned features in its de-
sign for this study. The pilot plant had a design capacity of
38 liters (10 gallons) of product water per minute. The plant
consisted of one cation exchange unit and one anion exchange
unit mounted together on a structural steel frame, as shown in
Figure 1. The units were of identical design, except for some
minor additions for lime slurry handling in the anion unit.
Each unit consisted basically of a loop made up of five func-
tional sections (treatment, backwash, pulse, regeneration, and
rinse sections) in four vessels (loading, overflow, pulse, and
regeneration vessels) as shown in Figure 2. The pneumatically
operated butterfly valves separated the vessels from one an-
other. All vessels were constructed of 15.2 cm (6 in) diameter
fiberglass column, except the loading vessel, which was a 20.3
cm (8 in) diameter fiberglass column.
The resins were transferred hydraulically around the en-
tire loop in the sequence of loading vessel, overflow vessel,
pulse vessel,and regeneration vessel. The frequency and quan-
tity of the resin transfer were regulated by a control panel.
Functions of Various Vessels
The specific function of each vessel in the pilot system
is summarized as follows:
Loading Vessel (Treatment Section)
The minerals in the form of cations and anions were re-
moved from the feed water in this vessel. The feed water en-
tered the vessel through distributor and left through collector.
The distributor and collector were designed to provide equal
distribution and collection of water over the cross-section of
the vessel. The distributor was a slot-opening type of design
without a screen, while the collector was provided with Johnson
well screen (approximately 0.4 mm opening) to prevent resin
from entering the product water.
Overflow Vessel (Backwash Section)
This vessel received the exhausted resin from the loading
vessel during the pulse period. The pulse water flowed out of
the top of this vessel, as indicated in Figure 2.
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Figure I. Photograph of CHEM-SEPS continuous counter-current
ion exchange pilot plant.
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TREATMENT
SECTION >
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SECTION
VALVE POSITIONS DURING CYCLES
RUN CYCLE PULSE CYCLE
VALVES OPEN VALVES CLOSED
VALVES OPEN VALVES CLOSED
A, WA, IN, EF, B,C, DANDPU. B, C, D AND PU. A, WA, IN, EF,
AND RE. AND RE.
Figure 2. Schematic diagram of CHEM-SEPS continuous
counter-current ion exchange contactor.
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During the normal operation period other than .the pulse
period, the backwash water entered this vessel and washed the
suspended solids and resin fines out of this vessel into the
resin recovery tank. The resin fines consisted of small resin
particles produced by mechanical attrition or osmotic shock.
A sight glass was installed in this vessel to allow visual
examination of the resin level. A proper resin level was main-
tained in the vessel to insure that a sufficient amount of resin
was in the unit for effective and accurate operation.
Pulse Vessel (Pulse Section)
The major function of this vessel was to initiate and con-
trol the movement of resin around the loop from one vessel to
the other. The cleaned but exhausted resin from the overflow
vessel was transferred to this vessel by gravity. The exact
amount of resin to be transferred was controlled by the upper-
limit ultra sonic sensor (Sensall), which sent out a signal to
close the valve as soon as the resin had reached the pre-set
level .
During the pulse period, high pressure water entered this
vessel and pushed the resin out of the vessel into the regenera-
tion vessel. The amount of resin movement was closely con-
trolled by the lower-limit ultra sonic sensor. The distance be-
tween the two sensors could be adjusted to provide the proper
amount of resin movement per pulse period. The actual resin
movement could be measured by means of the sight glass provided
in the vessel. A pulse timer was used as a backup for the
Sensall pulse control.
Regeneration Vessel (Regeneration and Rinse Sections)
In this vessel an appropriate amount of regenerant was
used to regenerate the exhausted resin, and the regenerated
resin was then rinsed free of excess regenerant before being
pulsed into the loading vessel. Teflon-coated Johnson well
screens were used to prevent the resin from entering the re-
generant distributor and regenerant waste collector. The waste
collector screen could be replaced easily when it became clogged
by chemical precipitate.
The product water was used as rinse water which entered
through the distributor and passed through the resin bed to
displace the excess regenerant. The rinse water was auto-
matically controlled by a conductivity probe inside the rinse
section. A flow meter was used to indicate the flow rate of
rinse water so that a proper flow rate could be maintained to
prevent any channeling through the resin bed.
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Resin Characteristics
Basically, the Chem-Seps continuous counter-current ion
exchange pilot plant was a modified combination system of a
strong acid cation exchanger and a weak base anion exchanger.
The cation column employed a mixture of strong and weak acid
cation exchange resins, instead of strong acid cation exchange
resin alone. By utilizing such a mixture of strong and weak
acid type resins, the excess acid, which was required for an
efficient regeneration of the strong acid resin, could be
utilized for regeneration of the weak acid resin because of its
higher proton affinity. The proper fractional amount of the
weak acid cation exchange resin in such a resin mixture was de-
termined by the characteristics of the feed water quality, more
specifically, the fractional amount of the bicarbonate ions in
the water to be treated. A combination of 50 percent strong
and 50 percent weak acid cation exchange resin mixture was
recommended by Chem-Seps for Pomona wastewater demineralization
study.
The resins selected for this study were Duolite C-20
strong acid cation resin, Duolite CC-3 weak acid cation resin
and Duolite ES-340 weak base anion resin. All these resins
were manufactured and supplied by Diamond Shamrock Chemical
Company, Redwood City, California. Some of the major physical
and chemical characteristics of the resins are shown in Table I
The cation resin mixture was regenerated with 2.5 percent
sulfuric acid solution, while the anion resin was regenerated
with either 1.5 percent calcium hydroxide solution (hydrated
lime slurry) or 4 percent ammonium hydroxide solution.
10
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SECTION 5
PILOT PLANT OPERATION
GENERAL
The installation of the Chem-Seps continuous counter-
current ion exchange pilot plant was completed in July, 1971 at
Pomona Research Facility. The initial start up of the pilot
plant operation was scheduled on August 2, 1971. However,
successful pilot plant operation was not achieved until March 1,
1972. During the unsuccessful trial run period, a series of
major system alterations and modifications were made on the pi-
lot plant to provide the necessary control mechanisms for the
process study.
The pilot plant was operated satisfactorily using
HpSO^ - NH/.OH mode of regeneration during the first half of
March, 1972. The operation was then converted to the
^504 - Ca(OH)2 (lime slurry) regeneration mode. An air strip-
ping tower was installed between the cation and anion units, as
illustrated in Figure 3. The C02 was stripped from the effluent
of the cation unit to prevent the formation of CaCOo precipitate
inside the anion unit. This operation mode was continued
through the first week of May, 1972. However, the operational
results, which will be discussed in Section 6, indicated that
the existing pilot plant system was not able to utilize lime
slurry successfully for regeneration. Consequently, the attempt
to achieve the utilization of the low cost lime slurry for the
regeneration of the weak base anion exchange resin was abandoned,
On May 16, 1972, the pilot plant operation reverted to the
H2S04 - NH^OH regeneration mode. The air stripping tower was
bypassed for this mode of operation. The pilot plant was oper-
ated continuously from 8:00 a.m. on Monday through 4:00 p.m. on
Friday every week throughout most of the investigation period.
The operation of the Chem-Seps continuous counter-current ion
exchange pilot plant was terminated on August 31, 1972, after
the expiration of the six-month minimum rental period of the
pilot plant.
OPERATING CONDITIONS
Some of the important operating conditions employed for
the pilot plant operation are described in the following under
each specific process function.
12
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Demineralization
The demineralization process was accomplished by the
treatment sections of both cation and anion units. Each treat-
ment section contained about 40 liters of resin between the in-
let and outlet levels, thus providing about one minute of
empty-bed detention time for a normal 38 1pm feed water flow.
The hydraulic loading in both treatment sections under the same
38 1pm flow was about 17.3 1/sec/sq m (25.5 gpm/sq ft). The
feed pressure was regulated at 3.5 kg/sq cm (50 psi) for the
cation unit and the residual pressure in the product water for
the anion unit was around 0.7 kg/sq cm (10 psi).
Two intermediate sampling points, F-l and F-2 , were pro-
vided in each treatment section as shown in Figure 4. Samples
could be taken from these points to evaluate the degree of the
resin exhaustion. Every 10 minute period the demineralization
cycle was interrupted by the resin pulsing operation. About 10
percent of the resin in each treatment section was replaced by
the freshly regenerated resin for every pulse cycle. The ex-
hausted resin was pulsed into the backwash section for backwash
process.
Backwash
Carbon-treated secondary effluent water, which was the
feed water for the pilot system, was employed as the backwash
water. The backwash flow rate was controlled by a throttle
valve and was indicated by the backwash flow meter. The flow
rate was set at 3.6 1/sec/sq m (5.2 gpm/sq ft) for both cation
and anion exchange resins. At this flow rate, the resin bed ex-
pansion was expected to be in the range of 50 percent to 100
percent5 depending on the actual process temperature and resin
type. Sufficient amount of freeboard was provided in each over-
flow vessel to accommodate the expected bed expansion.
The duration of backwash cycle was controlled by the same
programmer as the demineralization cycle, which was set for 10
minutes for each cycle. However, the 10 minute duration seemed
too long for the pilot plant backwash operation. A second con-
trol timer could be installed to reduce the duration of the
backwash cycle to save the backwash operation cost, but it was
not done in this study.
The suspended solids, resin fines and, possibly, some per-
fect resin beads in the backwash wastes of the cation and anion
units were collected respectively in the cation and anion resin
recovery tanks. Most of the perfect resin beads could be re-
covered from each resin recovery tank by flotation separation
method. Both suspended solids and resin fines should be washed
out of the cation and anion units to prevent plugging of the
collector screens and the resin beds.
14
-------
OVERFLOW
CM
CM
CM
CM
FEED
F-
\ \
8
\
RINSE
R-l a SRI
R-2 & SR2
RINSE TAIL-«-
CONDUCTIVITY
PROBE
1
(0
- PULSE
o
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(0
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30
d"=2.54cm)
F: SAMPLES FOR TREATMENT pH PROFILE
R= SAMPLES FOR RINSE pH PROFILE
SR: STRIPPED RESIN SAMPLES
S: SAMPLES FOR STRIP PROFILE
LR: LOADED RESIN SAMPLES
Figure 4. Layout of pilot plant sample points.
15
-------
Pulse
The purpose of pulsing the resin is to move exhausted
resin out of the treatment section and replace it with regen-
erated resin. Since the accuracy of the resin movement will
directly affect the pilot plant performance, the assurance of
the precision and reliability of the control mechanisms in the
pulsing operation is overwhelmingly important.
In the Chem-Seps pilot plant, the amount of resin move-
ment for every pulse cycle was controlled by two ultra sonic
sensors in each pulse section. The total resin movement was
regulated at 20.3 cm (8 in) per pulse cycle for both cation and
anion units, which was equivalent to about 3.7 liters (0.13 cu
ft) of resin per pulse. The rate of pulsing the resin was main-
tained in the range of 2.5 to 5.1 cm (1 to 2 in) per second and
was controlled by a manual throttling valve. The carbon-treated
secondary effluent pressurized at approximately 4.4 kg/sq cm
(62 psi) was used as pulsing water. The volume of water re-
quired to move the resin in every pulse cycle was practically
equal to the volume of the resin transferred.
The pulsing operation of the pilot plant was programmed to
be carried out one unit at a time, starting with the cation unit
first and then the anion unit. It usually took 10 to 15 seconds
for the cation unit to build up pressure to initiate the pulsing
and about 5 seconds to complete the resin transfer, while it
took from 20 to 40 seconds for the anion unit to initiate the
pulsing and about 5 to 8 seconds to complete the resin transfer,
Therefore, the total pulsing time ranged from 40 seconds to 68
seconds.
During the pulsing cycle, all other process functions
(demineralization, backwash, regeneration and rinse) were tem-
porarily suspended, and they resumed operations simultaneously
as soon as the pulsing operation was completed.
Regenerati on
Cation Exchange Resin
The regeneration of the cation exchange resin was accom-
plished by the sulfuric acid at a feed concentration of 2.5 per-
cent. A 4 percent sulfuric acid solution which was employed in
the previous study of a fixed-bed ion exchange system (2) was
found to cause serious plugging of the waste collector screen in
this study.
The flow of the 2.5 percent sulfuric acid was maintained
continuously through every ten minute run cycle at a rate of
approximately 38 liters (10 gallons) per hour. At this normal
flow rate, the actual detention time for the sulfuric acid in
16
-------
the regeneration section ranged from 20 minutes to 60 minutes,
depending on the rinse water flow conditions. The flow rate of
the regenerant was controlled and regulated by a throttling
valve and a flow indicator on the regenerant feed line.
The average feed pressure for the regenerant flow was about
1.0 kg/sq cm (15 psi) at the normal flow rate of 38 liters (10
gallons) per hour. However, the feed pressure was found to vary
from 0.6 to 1.4 kg/sq cm (8 to 20 psi) from one run cycle to an-
other, depending on the conditions of the packing of the resin
bed and the plugging of the waste collector screen inside the re-
generation section. This feed pressure variation caused sub-
stantial rate fluctuation for the regenerant flow.
Anion Exchange Resin
During the study period with the hLSO.-CatOHK regeneration
mode, a 1.5 percent Ca(OH)2 slurry was used as the regenerant for
the regeneration of the weak base anion exchange resin. The lime
slurry was initially made up from low cost slacked lime. How-
ever, it was found that the slacked lime contained some small
"chunks," which acted like resin particles and traveled co-
currently with the resin into the treatment section, thus de-
grading the quality of the product water. Consequently, finer
quality hydrated lime was used to replace the slacked lime for
making up the 1.5 percent Ca(OH)2 slurry after April 3, 1972.
The flow rate of the lime slurry was maintained approximately at
23 liters (6 gallons) per hour. At this flow rate, the deten-
tion time for the lime slurry in the regeneration section was
about 100 minutes, provided the rinse water was effectively dis-
charged through the rinse tail as shown in Figure 3.
The H2SO/i -Ca(OH)2 regeneration mode was converted to
H^SO^NH^H mode after two months of poor performance associated
with the lime slurry application. During the H2S04-NH40H oper-
ation period, the flow rate of the 4 percent NH^H was maintained
at about 7.6 liters (2 gallons) per hour under an average feed
pressure of 1.3 kg/sq cm (18 psi). Unlike the cation unit, the
feed pressure for the NH^OH in the anion unit was quite consis-
tent throughout the entire period of study. The actual deten-
tion time for the ammonium hydroxide solution in the regeneration
section was estimated to range from 30 minutes to 270 minutes,
depending on the rinse water flow conditions.
Rinse
The demineralized water was employed as the rinse water for
both cation and anion units. The flow of the rinse water was not
continuous through the ten minute run cycle. The actual duration
of the rinse water flow was controlled by the conductivity probe
installed in each rinse section. The rinse flow was turned off
automatically when the conductivity measured by the probe dropped
17
-------
below the pre-set level. The rinse flow was regulated by a
throttling valve and maintained at 7.6 liters (2 gallons) per
minute for both cation and anion units.
Since the duration of the rinse flow varied slightly from
cycle to cycle, the total volume of rinse water used also varied
somewhat. The average amounts of rinse water for cation and
anion units during the operation of the ^SC^-NH^OH regeneration
mode were equal to 3 and 4 bed volumes of the processed cation
and anion exchange resins, respectively.
For the H2SO/j-Ca(OH )£ regeneration mode, the rinsing con-
dition for the cation unit was similar to the aforementioned one.
However, the volume of the rinse water for the anion unit was
substantially increased to 12 bed volumes of resin. Therefore,
a rinse tail had to be installed in the anion unit to remove
this big stream of rinse water to prevent the unfavorable di-
lution of the regenerant flow,
SAMPLING AND TESTING
During the pilot plant operation period, one-hour composite
samples from the various streams of feed, product and brine were
taken daily between 7:30 a.m. and 8:30 a.m. for performance
evaluation. In addition to this sampling schedule, a conduc-
tivity probe was installed in the product line to record the en-
tire sequence of product quality. All samples were analyzed at
Pomona Research Laboratory for the concentrations of the major
cations and anions, total dissolved solids, total alkalinity,
acidity, pH and conductivity.
Besides the daily one-hour composite samples, some grab
samples were also frequently taken at various points, as shown
in Figure 4. The grab samples taken from the F-l and F-2 sam-
pling points were used to determine the pH loading profiles of
the treatment sections in both cation and anion units. The
efficiency of the rinsing operation was evaluated by the samples
taken from the R-l and R-2 sampling points. The resin samples,
SRI and SR2, were used for determining the extent of chemical
stripping or state of regeneration, while the degree of resin
exhaustion was determined by the LR resin samples. Some grab
samples were also taken at the S-l, S-2 and S-3 sampling points
for the evaluation of the rate of decreasing strength of the
regenerant along the path of the regenerant flow. All these
grab samples served to monitor the various process functions and
optimize the overall pilot plant performance.
The analytical procedures for the cations and anions were
specified in the Standard Methods for the Examination of Water
and Wastewater (4), while the exchange capacities of both ex-
hausted and regenerated resins were determined by the simplified
18
-------
methods as described in Appendix A. The analytical procedures
for determining the total salts in the samples of S-l , S-2, S-3
and regenerant waste were presented in Appendix B,
19
-------
SECTION 6
RESULTS AND DISCUSSIONS
During the first part of the study, the pilot plant opera-
tion was interrupted many times, with the shut-down period vary-
ing from several days to as much as three months. The inter-
ruptions were required to make necessary repairs and modifica-
tions of the system control mechanisms. As a result of this
on-and-off operation mode, most of the experimental data col-
lected before March, 1972 are not considered significant, and
thus they are not included in this report. No major mechanical
problem was encountered during the second part of the study from
March to August of 1972, which was officially considered as the
six month minimum rental contract period for the Chem-Seps pilot
pi ant.
The long delay in the formal start-up of the pilot plant
operation has resulted in a cancellation of the long-term study
of the continuous counter-current ion exchange process. There-
fore, the experimental results and discussions presented in this
report are based on the six month short-term study.
REMOVAL OF IONIC IMPURITIES
When the H^SC^-Ca (OH) 2 regeneration mode was employed, the
rinsing of the regenerated resins in the anion unit was never
satisfactorily achieved. Both calcium carbonate and calcium sul-
fate precipitates were transferred into the treatment section
with the anion exchange resin during the pulsing operation.
These precipitates were slightly dissolved in the water inside
the treatment section and caused serious contamination of the
product water. The high concentrations of the calcium and sul-
fate ions in the final product water, as indicated in Table 2,
fully demonstrated the contamination effect of the chemical
precipitates.
The concentration of the sulfate ion was also frequently
found to be higher in the cation column effluent than the feed
water during both series of operations using two different re-
generation modes. The increase of the sulfate ion in the cation
column effluent could result from the dissolution of the calcium
sulfate precipitates embedded in the cation exchange resins, or
it could result from an inadequate rinsing of the excessive sul-
furic acid out of the cation exchange resins. As a result of
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this additional loading of the sulfate ion, the removal effi-
ciencies of the chloride and orthophosphate ions were sub-
stantially reduced in the anion exchange column. Consequently,
the overall removal of the ionic impurities, expressed as IDS
removal, during the HoSQ^-CaCOH^ operation series was very
poor. The IDS removal was found to vary widely from 10 percent
to 70 percent, with an average of 50 percent.
The contamination of the final product water by the cal-
cium sulfate and the calcium carbonate precipitates in the
H^SO^-CatOH^ mode of operation was eliminated by the use of an
ammonium hydroxide solution, instead of lime slurry, as the re-
generant for the anion exchange resin. This improvement is
clearly demonstrated in Table 3 and Table 4. Table 3 shows the
best set of experimental results obtained during the h^SO^N^OH
operation series, while Table 4 is the summary of a one month
period steady-state operation under the identical operation con-
ditions which have been given in Section 5. This one month
period, which will be discussed later, occurred near the end of
the entire study. It was characterized by a minimum of mechani-
cal problems.
As indicated in Table 3, the concentration of the ammo-
nimum ion was increased in the anion column effluent, thus re-
ducing the removal efficiency from 86 percent as achieved by the
cation exchange column alone to 69 percent. This adverse effect
on the ammonium ion removal by the anion exchange resin could be
attributed to the inadequate rinsing operation in the anion col-
umn. Several other operating conditions, within the limitations
of the existing pilot plant, were tried to improve the rinsing
operation, however, all those alternatives failed to show any
improvement without deteriorating the overall removal of the
anions. As shown in Table 4, the ammonium ion removal during one
month of steady-state operation period was only 41 percent com-
pared to 69 percent removal during the best run. This might have
been complicated by the problems associated with the resin puls-
ing operation.
REGENERATION EFFICIENCIES
During the steady-state operation period, the average prod-
uct flow rate was about 34 1pm (9 gpm). Based on this average
flow rate and the average ion concentrations as shown in Table 4,
the chemical loadings for the cation and anion exchange resins
per minute of the run cycle were estimated to be 340 milli-
equivalents and 190 mi 11iequivalents , respectively. The bicar-
bonate ion was not included in the estimate for the anion ex-
change resin. In order to maximize the regenerant utilization,
only equivalent rates of sulfuric acid and ammonium hydroxide
dosages were applied to the respective regeneration sections to
match the chemical loadings.
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Under such stoichiometric regenerant feed rate for both
cation and anion exchange resins, the average removal rates by
the respective resins were found to be about 300 milliequiva-
lents per minute for cations and 170 mi 11iequivalents per min-
ute for anions. Therefore, the regeneration efficiencies as
calculated by the following equation were about 88 percent and
90 percent for the cation and anion exchange resins, respec-
ti vely.
Regeneration Efficiency (%)
Total equivalents of ions removed
Total equivalents of regenerant applied
X 100
The regeneration efficiencies achieved by this study were
slightly higher than those of the two-stage fixed-bed ion ex-
change process (2). However, the average TDS removal demon-
strated by this pilot plant operation was only about 82 percent,
which was lower than the 88 percent achieved by the two-stage
fixed-bed ion exchange process at similar regenerant dosage
level. Furthermore, the regeneration efficiencies of this pilot
plant study were not as consistent as those from the fixed-bed
system on a day-to-day basis.
During this regeneration operation, the most serious prob-
lem was the plugging of the waste collector screen and the waste
discharge line of the cation exchange column. The plugging prob-
lem was minimized by replacing the waste collector screen once
every two to three weeks.
BRINE CHARACTERISTICS
In this continuous counter-current ion exchange pilot plant
study, the rinse water and the spent regenerant were combined in-
to the same brine stream in each regeneration vessel. Therefore,
the brine flow rate for the cation and anion exchange columns
could be measured separately in each corresponding brine holding
tank. The average rate of brine flow for the cation column was
about 98 liters (26 gallons) per hour, while the average rate for
the anion column was about 72 liters (19 gallons) per hour. All
these rates were measured during the H2SOit-NH!+OH steady-state
operation period. The brine/product ratio for each brine stream
was estimated to be 4.8 percent and 3.5 percent for the cation
and anion columns, respectively. Thus, the combined brine/pro-
duct ratio for the continuous counter-current system operation
was approximately 8.3 percent which was slightly lower than the
11 percent required by the two-stage fixed-bed ion exchange pro-
cess. However, when lime slurry was used instead of ammonium
hydroxide, the brine/product ratio for the anion column alone
was found to be as high as 12 percent.
25
-------
Table 5 shows the average results of the chemical analyses
of the brine samples taken during the HaSCK-NH^OH steady-state
operation period. The strength of the brine produced in this
series of operations was practically equal to the two-stage
fixed-bed operation.
RESIN STABILITY
Before the initiation of the series of intensive pilot
plant operations on March 1, 1972, some resin samples from both
ion exchange columns were taken for screen analysis of the resin
particle size distributions. The typical results of these analy-
ses are presented in Table 6 for both cation and anion exchange
resins. The same type of analysis was applied to the resin
samples taken after the termination of the entire pilot plant
operations, and the typical results are also included in Table 6.
As indicated in this table, the particle size distributions for
both cation and anion exchange resins had been significantly
shifted toward smaller sizes as a result of the intensive plant
operations. However, most of the resin particles examined under
magnifier were still in perfect whole bead condition after the
entire series of operations.
Some screen analyses had also been performed on the resin
samples taken from the resin recovery tanks. The typical re-
sults are shown in Table 7. The anion resin recovery tank
seemed to have more fine resin particles as indicated in Table
7. However, it was found that fewer of the anion exchange resin
beads were chipped by attrition than the cation exchange resins.
About three percent of the cation resin beads in the resin re-
covery tank were chipped.
A complete survey of the resin inventory in each ion ex-
change column was made before the start of the intensive opera-
tion schedule on March 1, 1972, and the same survey was con-
ducted again on August 31, 1972 after the termination of the
pilot plant study. During this period, a total of 2760 hours of
onstream operations were conducted with the pilot plant. The re-
sults of the resin inventory survey indicated that about 1.5 per-
cent of cation resins and 4.5 percent of anion resins were lost
during the 2760 hours of operations. Most of the resin losses
might be attributed to the attrition effects of the various
valve operations along the loop path of the resin flow. The
resin fines thus produced were then stripped out of the column
by the backwash water.
During the pilot plant operations, no permanent impairment
of the resin operational exchange capacity was detected. There-
fore, it appears quite conservative to assume an annual 10 per-
cent replacement for cation resin and 20 percent for anion resin
in the process cost estimate.
26
-------
TABLE 5. AVERAGE BRINE CHARACTERISTICS DURING STEADY-STATE
PILOT PLANT OPERATION (HSO-NHOH REGENERATON MODE)
Parameter
Cal ci urn
Magnesi um
Sodium
Potassi um
Ammonuum (NH^-N)
Sul fate
Chi oride
Nitrate (N03-N)
Orthophosphate (PCM
Total Alkalinity (CaCOs)
Acidity (CaC03)
Total Dissolved Solids
PH
Conduct! vi ty
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
ymhos/cm
Cation
Col umn
Brine
1 ,000
258
1 ,970
185
302
9,290
62
0.32
8.6
570
11 ,700
3.4
11 ,400
A n i o n
Col umn
Brine
26
8.
54
6.
2,830
4,450
2,110
3.
344
2,930
7,240
7.
19,000
3
0
9
6
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Q.
N
-o
^
4->
oo
S-
,
o
3
4_)
oo
cu
S-
o
M_
CU
CQ
'a?
^ >
cu
^ .,-
'o *"
0 C
c °
_ -o
4-» ^
o *"""
1 \
^^ QJ
"^ -M
c
cu
""" (J
cu s~
01 a?
a
3
4_>
OO
s-
^
sC
>,
oo
cu
S-
0
cu
QQ
"S *-
fO ^
^_) -r
(/^
QJ
QJ
00 O 00 0
«^- CO i
S3- CM LO 1-^ i
i LO CM
CO CM O i i
CM IT) CM
CM IO t^. CM CM
CO
-------
TABLE 7. PARTICLE SIZE DISTRIBUTIONS OF THE RESINS FROM RESIN
RECOVERY TANKS
U.S. Standard Cation Resin Anion Resin
Sieve Number (% Retained on Sieve) (% Retained on Sieve)
20 14 3
30 41 15
40 32 35
60 10 32
80 1 10
29
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PROCESS RELIABILITY
The performance of the Chem-Seps ion exchange pilot plant,
as illustrated in Figure 5, was very inconsistent on a day-to-day
basis. Most of the problems encountered during the pilot plant
operations were mechanical in nature. The inaccuracy of the
resin pulsing control caused most serious damage to the process
performance. Other mechanical problems as listed below could
also be responsible for the inconsistency of the process per-
formance.
A. The conductivity control mechanism in the rinse
section was not sensitive enough;
B. The collector screen in cation regeneration
section was frequently plugged;
C. The sequence for the applications of the re-
generants and the rinse water was not
efficiently programmed; and
D. The feed pressure of the pulsing water might
not be high enough.
If all the above problems could be properly adjusted and
corrected, the performance of the continuous counter-current ion
exchange pilot plant would be more reliable and satisfactory.
During the last month of the pilot plant operation, the
plant was able to function steadily with relatively fewer
mechanical problems. The performance results of this one-month
period steady-state operations are shown in Figure 5 and Table 4.
Most of the process evaluations and the process cost estimates
have been based on the data collected during this relatively
trouble free steady-state operation period.
30
-------
640
600
560
520
480
440
o>
E
400
en
Q 360
<" 320
Q
> 280
240
200
O 160
120
80
40
0
I T I I I T
INFLUENT TDS
EFFLUENT TDS
-PERIOD OF^H
STEADY-STATE
OPERATION
10 15 20 25 30 35 40 45 50 55 60 65
OPERATION DAY (611172 -8l3ll72)
Figure 5. Performance of the CHEM-SEPS continuous counter-current
ion exchange system.
31
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SECTION 7
PROCESS COST ESTIMATE
A preliminary cost estimate, based on the Pomona pilot
plant operations, has been prepared for a 37,850 cu m/day (10
MGD) continuous counter-current ion exchange plant for waste-
water demineralization. The itemized cost estimates as well
as the assumptions used for making the cost estimates are shown
together in Table 8. As indicated in Table 8, the total pro-
cess cost (excluding the costs for carbon adsorption pretreat-
ment and brine disposal) for achieving an 82 percent deminerali-
zation is about 4.9^/1,000 liters (18.5^/1,000 gallons). This
cost will be increased to 7.4<£/l,000 liters (28.1^/1,000
gallons) if the carbon pretreatment cost is included.
In most wastewater reuse applications, it may require the
removal of an amount equivalent to that added during one domes-
tic use of the water, instead of requiring the removal of 82 per-
cent of the influent TDS as achieved by the Chem-Seps pilot
plant operation. In most cases, the amount added during one
domestic use of the water is about 300 mg/1 TDS. This is about
one-half of the influent TDS in Pomona wastewater. Therefore,
it is reasonable assumption that about one-half of influent TDS
will require removal in wastewater demineralization. It is
presently envisioned that product water of this quality, with
one-half TDS reduction, will be achieved by blending the high-
ly demineralized ion exchange product water with wastewater
which has been treated by carbon adsorption but has not been de-
mineralized. The total process cost of such a blended product
water in a complete recycle system is estimated to be about 5.2
-------
TABLE 8. PROCESS COST ESTIMATE
Plant Size = 37,850 cu m/day (day MGD)
£/1 .000 liters 1/1 .000 gallons
Capital
$1,200,000 0.8 2.9
Operation and Maintenance
Regeneration Chemicals
Sulfuric Acid 2.2 8.3
Ammonium Hydroxide 0.9 3.5
Resin Replacement 0.2 0.7
Power 0.3 1.0
Labor 0.3 1.3
MaintenanceMaterials 0.2 0.8
Total Process Cost 4.9 18.5
Assumptions:
1. Influent TDS = 600 mg/1; TDS Removal = 82%.
2. Regeneration Efficiencies: Cation = 88%; Anion - 90%.
3. Resin Replacement: Cation = 10%/year; Anion = 20%/year.
4. Sulfuric Acid = $36/ton; Ammonium Hydroxide = $80/ton.
5. Resin Costs: Cation = $!/£; Anion = $2/£.
6. Power = l<£/kwh.
7. Labor = 4 at $9,600/year.
8. Costs for carbon pretreatment (2. 5<£/l ,000 liters) and brine
disposal are not included in the estimate.
9. The capital cost is amortized for 20 years at 6% interest.
10. The estimate is based on August, 1973 material and con-
struction costs.
33
-------
total process cost can be reduced to 4.7<£/l,000 liters (18.1 <
1,000 gallons), including the cost for carbon adsorption pre
treatment but not the brine disposal.
34
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REFERENCES
1. Dryderi, F. D. , "Mineral Removal by Ion Exchange, Reverse
Osmosis and Electrodialysis," presented at the Workshop on
Wastewater Reclamation and Reuse, South Lake Tahoe,
California, June, 1970
2. Parkhurst, J. D., Chen, C., Carry, C. W., and Masse, A. N.,
"Demineralization of Wastewater by Ion Exchange," presented
at the 5th International Water Pollution Research Conference,
San Francisco, California, July, 1970.
3. Duolite Tech Sheet 103. 125 and 340, Resinous Products
Division/Diamond Shamrock Chemical Co./1901 Spring Street,
Redwood City, California.
4. Standard Methods for the Examination of Water and Wastewater,
APHA, AWWA. and WPCF. 13th Edition, 1971.
5. Chen, C., and Miele, R. P., "Wastewater Demineralization by
Two-Stage Fixed-Bed Ion Exchange Process," Final Report pre-
pared for Office of Research and Development, U.S. EPA,
September, 1977.
6. Higgins, I. R., Chemical Separations Corporation, Oak Ridge,
Tennessee. Private communication.
35
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APPENDIX A
ANALYTICAL PROCEDURES FOR DETERMINING
THE RESIN CAPACITY (6)
The weak base anion resin capacity, loaded (LR) or
stripped (SR), is determined by feeding an excess of a known
amount of standardized acid. The OH-form resin capacity (as
meq/ml or eq/1) is then obtained by dividing the acid sorbed
(meg) by the resin used (ml).
Example:
1. 25 ml of resin is used in a column.
2. 150 ml of 0.5 N_ acid is fed to the column. This
is equivalent to 75 meq.
3. Rinse the column with 3 to 5 bed volumes of dis-
tilled water.
4. Measure the volume of rinsed water and titrate
the acidity. Assuming 50 meq acidity is titrated.
5. The acid sorbed is then equal to 25 meq (75 meq
minus 50 meq).
6. The OH-form capacity is thus equal to 1 eq/1 or
1 meq/ml (25 meq/25 ml).
The same technique may be used to measure the H-form
capacity of strong acid type resin using standardized base.
However, it may not be used for weak acid type resin because
the resin capacity is a function of pH. The capacity at pH
13 to 14, when using NaOH, would not be the same as that for
water at pH 7.
The other common method is to strip off all cations
using about 10 bed volumes of 4 M HC1 or HNOs and analyzing
for all the cations removed.
Example:
1. 25 ml of cation resin is used in a column.
2. 250 ml of 4 M HN03 strip solution is applied.
3. Results of cation analyses:
CaT
Mg
Na+
.
=
1
200
400
800
rog/
tug/
mg/
1
1
1
=
=
=
1 .
0.
0.
2
4
8
g/i
g/i
g/i
36
-------
K+ , = 100 mg/1 = 0,1 g/1
NHu = 200 mg/1 = 0,2 g/1
4, Total equivalents of cations:
Catt = 1-2/20 X 250/25 = 0.60 (eq/1)
Mg^ = 0.4/12 X 250/25 = 0.33 (eq/1)
Na = 0.8/23 X 250/25 = 0.35 (eq/1)
K+ = 0.1/39 X 250/25 = 0.03 (eq/1)
* 0.2/18 X 250/25 = 0.11 (eq/1)
Total cations = 1.42 (eq/1)
37
-------
MgCl2
NaNOs
HR
APPENDIX B
ANALYTICAL PROCEDURES FOR DETERMINING
THE TOTAL SALTS (6)
If a sample has a mixture of salts in acid, the free acid
is titrated first. Then the sample is run through a bed of
H-form resin and titrated again. Subtraction of the free acid
gives a measure of the total salts. If a sample is alkaline,
then the free base is titrated. When running through a bed of
H-form resin the free base is sorbed to form water and salts
converted to acid.
CaR
MgR
+
NaR
HR
CaR
MgCl2 MgR
+ HR -* +
NaN03 NaR
NaOH HR
Recommended sample size is as follows:
1 - 10 meg/1
10-100
100 - 2000
2000 - 10,000
HC1
HN03
plus free H2SOt+
HC1
HN03
100 ml
10 ml
1 ml
0.1 ml
Any laboratory sized column of resin may be used;
commonly 25 ml to 100 ml. Standard 20-50 mesh resin is
satisfactory, but 50-100 mesh is preferred because its flow
rate is self-moderating. Also, because of the smaller resin
size, the exchange zone is sharper. It is important that the
38
-------
exchange capacity of the resin is never exceeded. Usually it is
regenerated after 1/2 to 2/3 of the capacity is used up. Rela-
tive to technique it is important that all the acid produced
from each sample is completely rinsed out of the bed. This is
commonly checked with pH paper. After familiarity one will
find that 3 to 5 bed volumes of rinse is adequate.
After exhaustion, the bed must be quite thoroughly re-
generated. This may be done by using 10 bed volumes of 4 M HC1
or HMOs. Remember to rinse out all free acid. The resin
shrinks a little on treatment with strong solutions. It is
helpful to fluff up the bed and let it resettle before the final
rinse. Never do this between regeneration because it is
necessary that all loaded resin be on top and the bottom of the
bed always 100 percent H-form.
39
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/2-77-152
2.
3. RECIPIENT'S ACCESSIOONO.
4. TITLE AND SUBTITLE
Wastewater Demineralization by Continuous Counter-
Current Ion Exchange Process
5. REPORT DATE
September 1977 (Issuing Date)
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
Ching-lin Chen
Robert P. Miele
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
County Sanitation Districts of Los Angeles County
Whittier, California 90607
10. PROGRAM ELEMENT NO. Task
1BC611-C611B SOS #3 B.08
11. CONTRACT/GRANT NO.
14-12-150
12. SPONSORING AGENCY NAME AND ADDRESS
Municipal Environmental Research LaboratoryCin., 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
wastewater demineralization study employing a 38 1 pm (10 gpm) continuous
counter-current ion exchange pilot plant, manufactured by the Chemical Separations
Corporation, Oak Ridge, Tennessee, has been conducted at the County Sanitation
Districts, Pomona Research Facility, Pomona, California. Under steady state condi-
tions IDS removal of 82% was achieved with a feed TOS of 500-600 mg/1. Monovalent
cation leakage resulting from inadayne rinse reduced TDS removal below that
obtained with a 2-stage fixed bed process tested at the same site.
The continuous counter-current ion exchange pilot plant has demonstrated
a promising regeneration efficiency for both cation and anion exchangers. The
brine volume produced by the process was approximately 8 percent of the product
flow, thus yielding a 92 percent water recovery. The annual resin operation losses
were about 5 percent for the cation exchanger and 15 percent for the anion exchanger.
These high resin losses, however, account for less than 5 percent of the total
process cost.
A cost estimate for a 37,850 cu m/day (10 MGD) continuous counter-current
ion exchange plant based on Pomona pilot plant operating results has been made. The
estimated total process cost of 4.8C/1000 gallons) was based on the use of carbon
treated secondary effluent with an average TDS concentration of 600 mg/1 to produce
a product water with 82 percent reduction in TDS.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.IDENTIFIERS/OPEN ENDED TERMS C. COSATI Field/Group
Demineralizing
Desalting
Ion Exchang
Purification
Water Reclamation
Wastewater Renovation
Continuous Counter-(
Ion Exchange
Current 13 B
13. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (This Report)
Unclassified
21. NO. OF PAGES
48
20. SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
40
S. GOVERNMENT PRINTING OFFICE 1977-757-056/6524 Region No. 5-11
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