WATER POLLUTION CONTROL RESEARCH SERIES
ORD- 14010DYH 12/71
   Neutradesulfating
   Treatment Process
for Acid Mine Drainage
    I'.S. ENVIRONMENTAL AGENCY

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          WATER POLLUTION CONTROL RESEARCH SERIES
The Water Pollution Control Research Series describes the
results and progress in the control and abatement of pollution
in our Nation's waters.  They provide a central source of
information on the research;  development and demonstration
activities in the Environmental Protection Agency,  through
inhouse research and grants and contracts with Federal,
State, and local agencies, research institutions,  and
industrial organizations.

Inquiries pertaining to Water Pollution Control Research
Reports should be directed to the Chief, Publications Branch
(Water), Research Information Division, R&M,  Environmental
Protection Agency, Washington, D.C. 20^60.

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        Neutradesulfating Treatment Process

               for Acid Mine  Drainage
                           by
                 Catalytic,  Incorporated
            Philadelphia, Pennsylvania  19102
                         for  the

              Environmental Protection Agency
                  Program No. 14010 DYH
                  Contract No.  14-12-518

                       December  1971
For sale by the Superintendent of Documents, U.S. Government Printing Office, Washington, D.C. 20402 - Price $1.00

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                  EPA Review Notice
This report has been reviewed by the Environmental Protection
Agency and approved for publication.  Approval does not
signify that the contents necessarily reflect the views and
policies of the Environmental Protection Agency nor does
mention of trade names or commercial products constitute
endorsement or recommendation for use.
                         ii

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                         ABSTRACT

Acid mine drainage pollutes many of the  surface streams of Appalachia.
A process for the treatment and removal  of  the major pollutants has
been develooed.  Iron and aluminum are removed by precipitation upon
addition of controlled quantities  of a neutralizer, such as sodium
bicarbonate.  Subsequently, a cyclic process  is employed in which
barium is eluted from a cation exchange  resin, and reacts with the
sulfate in the water to form a precipitate.   Barium is continuously
recovered and converted to a form  which  can be used to reload the
partially exhausted resin.  Sodium bicarbonate is also produced for
use in the neutralization reaction.   The only pollutant from the
process, hydrooen sulfide, is converted  to  sulfur and is sold as a
by-product.  Treatment sections are included  to insure complete
recovery of barium from all outgoing streams.  The process minimizes
the cost of sludae and^waste disposal while offering a product water
which meets the highest interstate water quality standards.  The
treatment cost for a 1MM GPD Plant is $2.69 per thousand gallons of
treated water.  The project was terminated  at the end of Phase I due
to the hioh estimated cost of treatment.

This report was submitted in fulfillment of contract number 14-12-518
between the Federal Water Quality  Administration and Catalytic,
Incorporated.

Key Words:  Barium elution, resins,  anions, carbon dioxide pressure,
            ion exchange.
                                 ill

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                             CONTENTS
Section                                                        Page

  I         CONCLUSIONS                                        xi
  II        INTRODUCTION
  III       EXPERIMENTAL - NEUTRADESULFATING
               Pretreatment of AMD
                  Theory
                  Composition of Feedwater
                  Pretreatment Tests
               Desulfating of AMD
                  Introduction
                  Model
                  Development of a Pressurized
                     Desulfating System
                  Factors Influencing Barium Elution
                  Elution Rate Studies
                  Continuous Tests
               Filtration
  IV        ACKNOWLEDGMENTS                                    25


  V         REFERENCES                                         27
  VI        APPENDICES                                         29
               A.  Statistical Analysis of AMD
               B.  Pretreatment Tests
               C.  Batch Elution Tests
               D.  One Million GPD Plant

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                              FIGURES


No.                      Name                                   Page

  1    Block Diagram for Brine Desulfating Process                 3

  2    Equilibrium Between Barium Resin  and Solutions  of         17
      Sodium  Chloride  (IR-120F)  (*)

  3    Equilibrium Between Barium Resin  and Solutions, of         18
      Sodium  Chloride  (IR-118)  (**)

  4    Equilibrium Between Barium Resin  and Solutions  of         19
      Calcium Chloride (IR-120F) (*)

  5    Equilibrium Between Barium Resin  and Solutions  of         20
      Calcium Chloride (IR-118)  (**)

  6    Iron and Aluminum in  Solution after Neutralization        48

  7    Bicarbonate and  pH as Functions of Soda Ash Addition      49

  8    Typical Pretreatment  Run  for Iron Rich  Solutions           50

  9    AMD Pretreatment: pH  as a  Function of Mixing Time         51

 10    Settling Rates of Synthetic AMD Solutions                  52

 11    Settling Rates of Synthetic AMD Solutions                  53

 12    Effect  of  Calcium Chloride Concentrations  on Elution      55
      of Barium  for IR-120F (*)

 13    Effect  of  Sodium Chloride  Concentrations on Elution       56
      of Barium  for IR-120F (*)

 14    Effect  of  Calcium Chloride Concentrations  on Elution      57
      of Barium  for IR-118  (**)
                          I
 15    Effect  of  Sodium Chloride  Concentrations on Elution       58
      of Barium  for IR-118  (**)
 (*)   IR-120F  is  Rohm & Haas  trademark for  strong  acid cation
      exchange resin of the  gel  type with fine  -40 to  +50
      particle mesh and 8% cross-linked.

(**)   IR-118 is a similar  resin  with -14 to +50 mesh particle
      assay (commercial grade) and 4%  cross-linked.
                                 vi

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                                  FIGURES


No.                       Name                                  Page

16     Continuous Ion Exchange Desulfating Apparatus             59

17     Elution of Barium from IR-120F for Solutions of           60
       Sodium Chloride

18     Elution of Barium from IR-120F for Solutions of           61
       Calcium Chloride

19     Elution of Barium from IR-118 for Solutions of            62
       Sodium Chloride

20     Elution of Barium from IR-118 for Solutions of            63
       Calcium Chloride

21     Effect of Anion on Elution of Barium from IR-Resins       64

22     Block Diagram for AMD Neutradesulfating                   69

23     Effect of Plant Capacity on the Unit Cost of              92
       Neutradesulfating

24     Effect of Capital Investment on the Unit Cost of          94
       Neutradesulfating
                                 vii

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                           TABLES


No.                      Name                             Page

I      Volume Occupied by Pretreated AMD Sludges              9

II     Norton Feed Water Profile from Statistical
       Tabulations                                           13

III    Batch Elution of Barium by Brine Solutions            15

IV     Continuous Ion Exchange Desulfating Runs              23

V      Statistical Analysis of AMD, Norton Site:
       Composite (*)                                         33

VI     Statistical Analysis of AMD, Norton Site:
       Stream G-lA                                           34

VII    Statistical Analysis of AMD, Norton Site:
       Stream G-l                                            35

VIII   Statistical Analysis of AMD, Hollywood Site:
       Proctor 1                                             36

IX     Statistical Analysis of AMD, Hollywood Site:
       Proctor 2                                             37

X      Statistical Analysis of AMD, Hollywood Site:
       Tyler Run                                             38

XI     Statistical Analysis of AMD, Hollywood Site:
       Bennett's Branch                                      39

XII    Statistical Analysis of AMD, Hollywood Site:
       Stream Combination No. 1                              40

XI\J   Statistical Analysis of AMD, Hollywood Site:
   v    Stream Combination No. 2                              41

XIV    Statistical Analysis of AMD, Hollywood Site:
       Composite                                             42

XV     Summary of Pretreatment Tests                         44

XVI    Sludge Settling Tests                                 47

 (*) Tables  V to XIV are statistical tabulations  from available
    data  on acid mine drainage samples  at  the considered sites,
                             VI11

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                              TABLES


_No.                          Name                              Pa

XVII   Production Plant Daily Traffic Schedule                  78

XVIII  Production Plant Utilities Summary                       79

XIX    Capital Investment for Neutradesulfating Plants          86

XX     Economic Evaluation for a 1MM GPD Neutradesulfating      87
       Plant
                               ix

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                              DRAWINGS
 No.                          Title                             Page

R-211    Engineering Diagram of a 1MM GPD Production Plant       95
         for Neutradesulfating of Acid Mine Drainage

R-212    Engineering Diagram for the Waste Recovery Section      99
         of a 1MM GPD Neutradesulfating Plant

B-211    1MM GPD Production Plant Layout (Sketch)               102

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This renort concerns the results  of an  investioation to develop a process
for the treatment of acid mine drainaae.   The work was based on the sta-
tistical data which was developed for mine drainage samples at several
sites at Norton, West Virainia and Hollywood, Pennsylvania.  The process
was intended for river v/aters containing  iron and aluminum pollutants in
the range of 100-1000 ppm, and sulfates in the  1000 porn range, but is
readily adaptable to waters over  a wide range of compositions.

Laboratory data were presented in the report for the neutralization and
precipitation of iron and aluminum using  soda ash.  The bicarbonate forn
was shown to work equally well.  Variables studied include mixing and
aeration time, pH, and amount of  neutralizer added.  Settling of the sludne
was shown to be of the uniform suspension (blanket settling) tyne and
typical curves are presented.  Equilibrium data from batch testing were
presented for the elution of barium from  ion exchange resins in the .001
to  .01 N sodium and calcium salt  ranges;  the selectivity ranged from 2 to 1
in  favor of the calcium ion.

In  the process which v/as developed, an  ion exchange cycle is utilized to
continuously remove sulfates from solution.  Laboratory evidence, using
a continuous bench scale apparatus, for desulfating of synthetic AMD so-
lutions shows that for effective  desulfating, the solution must contain
a certain quantity of anions other than sulfate to properly elute bariun
from the resin and to cause precipitation of insoluble barium sulfate in
the solution.  A one to one ratio of anions such as chloride or bicar-
bonate to sulfate is needed for complete  desulfating.  Statistically, the
acid mine drainane which has been treated to remove iron and aluminum has
an  insufficient amount of "other" anions  in comparison to the sulfate
which it contains in order to achieve maximum desulfating efficiences.
Continuous laboratory desulfatinq runs  used sulfate bearing brines in the
statistical range of AMD found at the chosen sites.  Near 100% sulfate
removal was accomplished for solutions  which had an equivalent ratio of
1:1, ClrSOd.  However, for solutions with increasing amounts of sulfate
over chloride ions, the desulfating efficiency  decreased rapidly.  By
adjustments of the feed snlits and the  use of fresh resin feed to the first
two exchanaers, a 65% efficiency  v/as attained for a 2:1 ratio solution.

Further manipulation of the system has  a  potential for high efficiencies
even with the lower anionic concentrations of the AMD.  The full develop-
ment of the ion exchanqe model would be required to determine the practical
limitations of the system.

Laboratory results show that the  process  would  be capable of substantially
reducina the concentrations of the major  pollutants in acid mine waters.
Virtually complete removal of iron and  aluminum to a pH of 6.0 can be ac-
complished in the neutralization  step.  Use of  large clarifier area alonn
with settling tube modules in the top 2 to 4 feet of wetted wall should
minimize shortcircuitina and carryover  of ironaluminum turbidity. Use
                                    xi

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of polyelectrolyte may be  necessary  for charqe neutralization of com-
plexed (hydrated iron and  aluminum)  suspensions.  Desulfating efficien-
cies in the exchanae treatment  cycle should be well above 50%, with
hioher ones expected with  the optimization of resin-feed snlit adjust-
ments.  For maximum efficiencies  (90-1002), a greater concentration of
anions other than sulfate  are needed than is present in AMD.

The process offers several  advantaqes such as a minimum of sludqe disposal
problems, low operating labor cost and a linht traffic schedule due to the
low chemical consumption requirements.  Almost all of the chemicals pro-
duced are reused in the process.  The only real disnosal problem is the
iron-aluminum sludoe from the pre-treatment-neutralization, which must be
deep-welled, or further dewatered for landfill, or trucked for reclamation.
The pollution effects for the process are nil, since hydrogen sulfide which
is formed in the carbonation reaction is converted to by-product sulfur.
Also, since barium in effluent  streams is a potential contaminant, snecial
treatment sections are desioned to provide minimum barium losses.  The
product water should be of sufficient quality to meet the hiqhest inter-
state stream standards.  All of the  barium sulfate which is produced will
be utilized in the process for  the regenerative needs of the resin.

A conceptualized desinn for the complete neutradesulfating process is
presented in the appendix.   Since it was recognized that the attractive-
ness of the process was due to  the hinh purity of the water which could
be produced, it was believed that a  method for increasing the anionic
concentration of the AMD was needed  for macimum desulfating efficiencies.
The method conceived was to utilize  high C02 partial pressures in the ion
exchangers under sealed conditions so that a temporary excess of ions
would be present due to the bicarbonate/carbonate, carbonic acid/C^
eauilibria.  No attempt has been  made to verify the method in laboratory
experimentation; the equipment  that  was conceived necessary for this
pressurization is included as an  option in the conceptualized process.

Based on the present projected  or scaled up technology, the unit cost of
a 1MM GPD capacity commerical unit would be $2.69/1000 gal. of feedwater
based on a 30-year payback period at 4.6%.  The total capital investment
was estimated at $4.96 MM dollars and represents about 35% of the unit
cost, or $.94/1000 gal.  The contract was terminated at the end of Phase I
due to the estimated hiqh  cost  of treatment.

Usinq a six-tenths factor,  the  prorated unit cost for a 10 MMOPD unit
would be $1.33/1000 gal.,  and $1.07/1000 gal. for a 25 M'i GPD plant, based
on 30 years at 4.6%.  If the capital cost of the above units could be re-
duced by 50%, the costs would be  $0.90/1000 gal., and $0.75/1000 gal.,
respectively, on the same  basis.  For the sized studied, increase in'the
size of the treatment plant does  not reduce the cost of treatment sin-
nificantly to make it competitive with other tynes of water treatment.

It is noted that fixed charqes  represent more than 50% of the treating
cost of a 1MM GPD plant.  If the  hinh equipment cost could be substantially
reduced through simplification  of the process, further reductions could
                                     XI1

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be effected in the direct and indirect costs which would result in sub-
stantially reduced treating costs.

The project was terminated at the end of Phase I due to the high estimated
cost of treatment.
Rec_o_minend_aj; i ons_

The economics of the neutradesulfating system application as studied and
piloted under bench scale conditions did not present a favorable projec-
tion for the statistical AMD water condition at Norton, West Virginia.
The design and construction of a 15,000 GPD pilot plant for this process
at this site is not recommended.

No further studies are recommended for this process on the particular sta-
tistical acid mine drainage water analysis presented for this contract based
on the high estimated cost of treatment.
                                  xiii

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                           INTRODUCTION
Discharge from active and abandoned mines are a major source of
river water pollution in the Appalachian Region of the United
States.  Five hundred billion gallons of mine drainage, containing-
5 to 10 million tons of acid, pollute over 10,000 miles of surface
streams and more than 15,000 acres of impounded waters annually (1).
About 40% of the mine drainage comes from active operations and the
remainder is from abandoned surface mines (25%), and shaft, drift,
and auger mines (35%).

The formation of acid mine drainage stems from water of surface
quality entering a mine and coming in contact with iron disulfide from
rock strata and coal seams.  In the presence of oxygen, ferrous sulfate
and sulfuric acid form, lowering the pH; this allows other compounds
containing Al, Mg, Mn, and Ca to be dissolved.  Further contact with
oxygen causes oxidation of ferrous to ferric sulfate and hydroxide,
both of which are soluble at the low pH (2-4).  In some instances, the
natural ecology of the land will partially neutralize discharges within
and around mine sites.  Therefore, since the nature and degree of pollu-
tion will vary greatly, the best method of treatment or control will
depend on the characteristics of the particular site.

Preventive measures, such as mining practices, reclamation, and control
of good water supplies, can be very helpful in holding mine drainage
from surface mines to a minimum.  However, preventive control will not
be entirely effective in every case, and for many abandoned mines, is
not applicable.  Source control methods are often not acceptable for
active operations.  In these cases, treatment of the mine drainage will
be necessary to eliminate the contaminates.  Therefore, to find solutions
to this problem, the Federal Water Quality Administration (FWQA) of the
U. S. Department of the Interior has begun a program which includes
Research and Development grants to industry.  These have lead to proposals
for the development of methods of acid mine drainage treatment.

In April, 1969 a research contract was awarded to Catalytic, Incorporated
by the FWQA to develop a process to reduce or eliminate the pollutional
effects of AMD.

It was thought that a modification of a process developed by the Bureau
of Mines (2) for the desulfating of seawater would have a potential to
treat AMD efficiently to produce a water meeting the highest interstate
stream water standards, at relatively low unit costs.

Project Breakdown

The project was divided into four phases:

               Phase I   -  Preliminary Studies & Design
               Phase II  -  Pilot Plant Design
               Phase III -  Pilot Plant Construction
               Phase IV  -  Pilot Plant Operation  (Optional)

                               -1-

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This report deals with work conducted under Phase I of the project.
The following were part of this effort:

1.  A laboratory study to define design criteria for the iron and
    aluminum removal (pretreatment) by neutralization and the desulfat-
    ing of AMD by use of cation exchange resins.

2.  A development of process flowsheets for both pilot and production
    units.

3.  Design and cost estimate of a small-scale pilot unit.

4.  Development of a conceptual plant design for a neutradesulfating
    plant of commercial size.

5.  An evaluation of the basic process feasibility.

6.  An economic study to determine the cost of producing clean water
    at various capacity and investment levels.

7.  A study to determine the marketability of BaSO^ waste.(*)

Basic Process

Part of the basic process is described in "Bureau of Mines" Report RI 6928
entitled "Preliminary Process Development Studies for Desulfating Great
Salt Lake Brines and Sea Water" by Darcy R. George, J. M. Riley, and Laird
Crocker (2).  The block diagram for this process is shown in Figure 1.

A modification of this process was the basis for Catalytic's work on a
75,000 gal./day desulfating plant to prepare sea water for feed to a
desalination plant at Wrightsville Beach, North Carolina (3).  The basic
steps to the complete modified process are described in R&D Progress
Report No. 289 entitled "Evaluation of Brine Desulfating Process as
Applied to Desalination - Phase I" (4).

However, the process had to be modified for acid mine waters which are
neutralized prior to desulfating, and which do not have the brine
concentrations present in seawater.  The new modification of the basic
process includes the following steps:

1.  Neutralization of AMD feed by addition of soda ash or bicarbonate
    to form insoluble iron and aluminum oxides and hydroxides which
    are removed by gravity settling.

2.  Contact of the "deironized" mine drainage with the barium form of
    a cation exchange resin.  This removes sulfate and carbonate ions
    as insoluble barium salts.  This may include a C02 or flue gas
    pressurization.
(*) Since all BaSO^ is used in the process, no such study was effected.

                               -2-

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                       BARIUM SULFATE CONVERSION

                  Noturol gos	

                                              Makeup BaS04
                    Cool
             r
               i
BaS04 coke      |

               U 
      Residue to waste
                              Roasting
                                                   Wttr
                              Ltaching
                              Filtration
                          9o(OH)g . Bo(HS)g
                              solution
  ION EXCHANGE AND SULFATE PRECIPITATION
                                                   SODIUM CARBONATE AND
                                                   SULFUR RECOVERY
         Barium loading
                                 NoOH-NoHS solution
Barium resin
       Barium elution and
       sulfote precipitation
                                 Sodium resin
                                                Wosti
                                                CO,
                                                       Carbonation
     BoSO4 and
     brine slurry
                                   1
          Row brine feed
._____IU~~
      Thickening and filtration
          Bo SQ,
           coke
                     I
                                                           NaHCOs
                                                                    HS
                                                               To sulfur or
                                                               H,SO
                                                               conversion
                                                   To NatCOj
                                                         recovery
         To magnesium and potassium recovery

                        1. - Block  Diagram for Brine Desulfating Process,

                                    -3-

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 3.   Separation  of the  desulfated,  neutral  mine  drainage and precipitated
     barium salts  from  the  resin.

 4.   Decarbonation and  complete  removal  of  barium from solution in the
     neutral,  desulfated mine  drainage,  followed by solids  separation,
     and final polishing of the  product  to  eliminate carryover turbidity.

 5.   Conversion  of calcium  and magnesium forms  in the partially exhausted
     resin to  the  sodium form  by contact with a  brine solution.

 6.   Conversion  of precipitated  barium salts  to  a soluble form by an
     atmospheric,  reduction roasting with carbon (pulverized coal)
     mixture.

 7.   Leaching  of the calcine to  form a barium salt solution.

 8.   Regeneration  of the partially  exhausted  sodium form of the resin
     with the  barium salt solution.   The effluent solution  contains a
     sodium salt (sulfide).

 9.   Conversion  of the  sodium  sulfide to bicarbonate by carbonation with
     C02 containing flue gas.  Hydrogen  sulfide  is evolved  and leaves in
     the flue.
10
     Recovery of hydrogen sulfide by conversion to elemental sulfur.

The complete modified process block diagram is presented with the Process
Description of the 1MM GPD conceptualized unit.

Steps 1, 2, 3, and 8 were carried out in bench scale operations which were
conducted at the laboratories of Cyrus William Rice Co. of Pittsburgh under
an FWQA authorized subcontract.

The ion exchange system had to be tested and adapted specifically for the
range of dilute solutions prevalent in the AMD at the chosen sites in
Norton, West Virginia, and Hollywood, Pennsylvania.  This system includes
the novel use of a partially exhausted resin and dilute operating ranges
for the resin regeneration and carbonation sections.

An alternate pilot plant design case is also presented which would contain
necessary equipment for the optional use of a pressurized ion exchange
system  (*), and is included as part of the conceptual design of the
commercial size unit.

The overall effect of these modifications is to increase the capital in-
vestment over that which would be expected in desulfating a brine con-
taining seawater.  Contributing factors include extra equipment, greater
capacities, and special materials of construction required for AMD service.
There is a resultant shift in the unit cost.


(*) A pressurized unit is theoretically capable of higher desulfating
    efficiencies, but as yet has not been laboratory tested.

                                -4-

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FWQA Authorization and Funding

The primary purpose of this FWQA contract is to develop an economically
and technically feasible process that will reduce or eliminate the
pollutional effects of acid mine waters.

The secondary purpose is to produce an effluent water that can be safely
discharged into the receiving stream, which may be used by local indus-
tries or municipal water systems (5).

The Federal Water Quality Administration initiated this study with the
award of a contract 14-12-518 to Catalytic on April 4, 1969.

The project was terminated at the end of Phase I due to the high estimated
cost of treatment.

Ground Rules for the Study

1.  The Pilot Plant shall be capable of handling 15,000 GPD of acid
    mine drainage.  The conceptual commercial unit shall be of suf-
    ficient size to treat 1MM GPD.

2.  The product water shall be of either interstate stream, potable
    or industrial quality.

3.  A site selection and evaluation shall be made.

4.  Product Water shall have a value of $0.22/1000 gal.

5.  By-product values are to be taken from the current market prices.

6.  The cost estimate will not include in-take facilities or sludge
    upgrading equipment.

7.  Cost of capital is to be 4.67 for 30 years.

8.  The cost of land will not be included.

9.  The dollar value is to be $.0075/KWH for electricity.

The composition profile of Acid Mine Drainage which was originally
suggested was not used as a basis for the design.  Rather, statistical
analyses were run for available stream data at Norton, West Virginia,
and Hollywood, Pennsylvania sites.  These are listed in Appendix A.
The statistical mean for a composite of the two Norton sites was the
basis for the designs presented in this report.
                               -5-

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Preliminary Studies and Design

A 1MM GPD plant preliminary design and economic evaluation is presented
in Appendix D of this report.  The design of a 15,000 GPD desulfating
pilot plant was also completed.  The unit would include only sections
which are vital in the verification of the technical feasibility of a
barium loaded ion exchange resin cycle in the treatment of neutral
(iron free) mine drainage feed waters.  A portable (skid-mounted) design
was presented which includes a description of the proposed process and
equipment, flowsheets, layout plan, and material balance, along with
capital cost estimates for several variations.  The FWQA has made this
a separate document which will be filed in their office in Washington.
                                -6-

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                  EXPERIMENTAL - NEUTRADESULFATING
A laboratory study was conducted during Phase I of the contract to
establish a basis of design for both the pretreatment and desulfating
of acid mine drainage.  Pretreatment studies were centered around
treating synthetic mean Norton and Hollywood compositions with sodium
carbonate.  Effects of aeration and mixing time on pH were determined.
Settling tests were conducted for the resulting sludge precipitates.
Also tests were conducted in order to determine the efficiency of
Neptune Microfloc(*) tubes in improving the settling characteristics
of the sludge precipitate.

Desulfating tests were fashioned after those described by George,
Riley and Crocker (2).  The purpose was to determine equilibrium
conditions for dilute solutions of sodium and calcium chlorides (0.1
normal and less) in batch elution runs, and to determine desulfating
efficiencies of comparable solutions of sulfate bearing brines in a
continuous four-stage ion exchange cycle.  These tests were to study
the feasibility of desulfating Norton feed water from mean down to
lean Norton conditions using barium loaded resins.

Pretreatment of AMD

Theory

Iron (III) and aluminum are removed from acid mine drainage by precip-
itation as ferric oxide and aluminum hydroxide when the AMD is reacted
with sodium carbonate or bicarbonate.  There are two steps to the
reaction:

        Neutralization:
        3 Na2C03 + M2(S04)3(**)   	*   M2(C03)3 + 3 Na2S04      (I)

        Decomposition:
        M2(C03)3 + xH20           	   M20  . xH20 + 3 C02

        M2(C03) + (3+x) H20       	   2 M(OH)3 .  xH20 + 3 C02  (II)

or if bicarbonate is used (alternate):

        Neutralization:
        6 NaHC03 + M2(S04)3       	   2M(HC03)3 + 3 Na
(*)  Mention of commercial products does not imply endorsement by the
     Federal Water Quality Administration.

(**) M = Mineral

                                -7-

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        Decomposition:
        2 M(HC03)3              	    2 M(OH)3 + 6 C02    (HA)

The C02 that is formed during the decomposition step will react with
the available carbonate and the following reversible reaction occurs:

        Na2C03 + C02 + H20      -g	*>    2NaHC03             (III)

The equilibrium ratio of bicarbonate to carbonate is a function of the
COo partial pressure, sodium normality, the solubility of CC^, and the
temperature.  The final pH of the solution depends on this ratio.

If some of the iron is in the ferrous form it must be sparged with air
so that insoluble ferric hydroxide will be formed.  The oxidation reaction
is as follows:

        4 Fe(OH)2 + 6H20 + 02	>    4 Fe(OH)3 + 4H20    (IV)

This indicates that one pound of air is needed to oxidize 2.5 Ibs of
ferrous hydroxide; however, a large excess of air is required to bring
about a reasonable rate.

Composition of the Feedwater

Statistical analyses were conducted to determine the range of ionic con-
centrations which would be present in AMD at the Hollywood, Pennsylvania,
and Norton, West Virginia sites.  These may be found in Appendix A
(Tables V to XIV).

Since most of the iron is in the ferric form at the Norton site, aeration
to oxidize ferrous to ferric iron is of secondary importance.  However,
at the Hollywood site, aeration time and rate are critical since the
ferrous form predominates.

Pretreatment Tests
For most of the tests which were conducted, synthetic acid mine drainage
solutions were composed of iron solutions in the 1000 ppm and 100 ppm
range to resemble what was thought to be mean Hollywood and Norton feed-
water compositions, respectively.  However, statistical data later
received from the Hollywood site in general, were much lower in iron
concentration than was originally anticipated; Proctor Number 2 had a
mean iron (II) concentration of 850 ppm, and Proctor Number 1 had 250,
but all the others had iron concentrations below 100 ppm; thus, the
average value of the combined streams is expected to be closer to the
Norton mean of 120 ppm iron (III).

The tests were run by adding a measured amount of soda ash solution to
a 500 cc synthetic solution consisting of dissolved iron, aluminum, and
calcium sulfates, with sulfuric acid added in some cases to dissolve the
salts.

-------
Subsequently, the treated solutions were mixed for 1-2 hours, air
sparged for 1/2 hour, and then introduced to a 500 cc graduated
cylinder in which quiet settling was observed for 2 hours.

Appendix B, Table XV is a summary of the pretreatment tests, and
Table XVI shows the settling rates of the corresponding sludges.
These results show that the addition of near stoichiometric amounts
of soda ash result in an immediate rise of pH to between 5 and 6.
Thus, reactions I and II are instantaneous.  However, the reactions
rates for III and IV are much slower, and, in general, require an
hour for mixing and air sparging.

Some of the results are plotted in Appendix B.  Figure 6 shows aluminum
and iron (III) residual concentrations as a function of pH; and Figure
7 gives the bicarbonate concentration and pH as a function of the
amount of soda ash added.  Figure 8 is a typical curve for treatment of
iron rich solutions with soda ash.  From Figure 7, it is noticed that
less than a stoichiometric quantity of soda ash (90-96%) is required
for complete iron and aluminum removal.  This is due to precipitation
of some unreacted iron and aluminum sulfates at pH of 5 to 6.  However,
the iron concentration of Figure 6 increased rapidly as the pH decreases
below 6; therefore, pH is the limiting and controlling factor in soda
ash addition.

The sludge resulting from soda ash pretreatment of synthetic acid mine
drainage solutions consists of hydrated iron and aluminum oxides and
hydroxides, along with some calcium carbonates.

Tests were also run using the bicarbonate rather than soda ash; they
show only slight differences in the pretreatment curve and the same
overall effect will be realized.

The two-hour settling tests were run in a 500 cc graduated cylinder.
In the Appendix B, Figures 10 and 11 show the sludge volume measured
at the interface (cc.) as a function of time.  The iron rich solution
(1150 ppm) experienced strong spherical interferences resulting in the
hindered settling characteristics; the "lean" iron curve containing
120 ppm as in the mean Norton feedstream, was much steeper, thus requir-
ing a smaller thickener area.  Table I below shows the final dilution
ranges of typical AMD solutions.

                              TABLE I

              Volume Occupied by Pretreated AMD Sludges

              Iron (III)      y%(*)           W% Solids

               1150 ppm      20-35%           0.8-1.3%
                120 ppm      4-7%             0.8-1.1%

              (*) Based on final dilution volume.

-------
The conclusions drawn from the pretreatment testing were used to
establish the following criteria for the process design.

1.  Soda ash should be added in line to allow a solution of
    approximately 6.0 pH to enter the reactor.

2.  One hour retention time with vigorous agitation and simul-
    taneous pressurized air sparging in a back-mix reactor will be
    sufficient to oxidize ferrous iron, remove any excess C02,
    precipitate remaining iron and aluminum, and sustain a bicarbonate/
    carbonate residual environment which is favorable for the
    desulfating process.

3.  A requirement of 2400 square feet per ton dry solids per day
    along with 8 feet wetted wall, would provide for sufficient
    sludge thickening.

It is believed that net point charges on individual particles may
generate an electric field which would bring about the formation of
colloids, thus preventing agglomeration of the spheres due to these
interferences.  As the concentration increases these interferences
offer a greater hindrance to particle settling.  Use of a dispersant,
or a flocculent, could result in formation of large groups of neutral
clusters which rapidly settle.  Further investigations are warranted,
using a variety of industrial polyelectrolytes.

Desulfating of AMD

Introduction

Sulfate is removed from neutralized (pretreated) acid mine drainage
when the AMD is contacted with a strong acid cation exchange resin
which is loaded with barium.  The rate of desulfating is a function
of the rate at which barium can be eluted from the resin.  Studies
were conducted to determine elution rates and desulfating efficiencies
and the factors which affect the same.  These tests were of two types:

1.  Batch Elution Tests were run to determine equilibrium conditions
    for synthetic solutions of sodium and calcium salts in the 0.001
    to 0.1 normal range.

2.  Continuous desulfating runs were conducted to determine the
    factors involved in offsetting and preventing the attainment
    of equilibrium, and the degree to which it could be accomplished.
    By the use of these factors, a temporary environment of barium
    ions over the equilibrium concentration could be maintained in
    solution to bring about complete sulfate removal.
                                -10-

-------
Model

During the course of these tests, the following reaction model was
developed which served as a basis and guide for the study  (6).  When
a particle of barium loaded resin is placed in contact with an ionic
solution bearing sulfates, the following steps occur:

1.  Diffusion of cation  (other than barium) to the site (resin surface).

2.  Diffusion of same from the surface into the porous resin.

3.  Displacement (Elution) of barium by the cation according to the
    reactions:

              2M+ + BaR2   ,       2MR + Ha"1"1"     (V)

              M4"4" + BaR2  >+ '      MR2 + Ba*4"

              (NOTE: M = Alkali or Alkaline Earth Metal Ion)

4.  Barium diffusion through the pore solution to surface  of resin or
    desorption.

5.  Barium diffusing into the bulk solution.

6.  Barium reaction with sulfate ions:

             Ba44" + S04=  	  BaS04          (VI)

7-  Barium sulfate precipitation (crystallization)

8.  Barium sulfate crystal growth

The first five steps are reversible by nature and require  the existence
of comparable anions that will yield a soluble barium salt (e.g.: HC03~
or Cl  ).  If only sulfate anions were present in quantity,  then steps  4
and 5 would not occur and barium sulfate would form and precipitate  in
the resin pores, thus resulting in "blinding" of the resin.  This blind-
ing effect coupled with  the relative immobility of the "bulky" sulfate
ions results  in very low desulfating efficiencies.  Steps  1 to 5 can be
lumped as the equilibrium section, whereas steps 6, 7, and 8 are thermo-
dynamically favorable and irreversible and are the non-equilibrium part.
The former is the slower and therefore rate controlling part of this
heterogenious system.
                                -11-

-------
To insure that the formation of the solid barium sulfate product does
not terminate or inhibit its own propagation, the equivalents of dissolved
barium in the bulk solution must be greater than or equal to the sulfate.
Thus, desulfating efficiency would be high for feed solutions containing
an anion(*)/sulfate ratio of one or more.  However, the percent of
desulfating would decrease rapidly as the ratio dropped below 1.

Development of a Pressurized Desulfating System

Since the Norton feedwater profile in Table II indicates that the sulfate
in the "mean" water is about 0.019 normal, over 1100 ppm bicarbonate
would be required for a ratio of 1,  However, only 100 ppm is expected to
be present in mean Norton pretreated feedwater.

One solution is to add an additional quantity of sodium bicarbonate or
carbonate to the desulfating section feed stream.  However, this would
be costly due to high sulfuric acid requirements for decarbonation.  A
more economical approach is to use pressurization with CC>2 and create a
"temporary" 1;1 ratio in the barium exchanger train due to formation of
bicarbonate under high 62 partial pressures.  The reaction

            Na+ + H2C03  -y->-  NaHC03 + H+     (VII)

would therefore be thermodynamically favorable under such conditions.
The initial pressures would be high (75 psig) due to high initial sulfate
concentration, and would be gradually reduced until only about 3 psig
remains in the last exchanger due to previous partial desulfating.

One last possibility is to maintain C02 pressures from the pretreatment
sections, but this is thought to greatly hinder iron removal due to
reaction of the hydroxide with C02 to form iron bicarbonate which could
be favorable under high CO? partial pressures.

Further experimentation either in bench or pilot scale is needed to
verify the pressurized desulfating model.

Factors Influencing Barium Elution

Providing sufficient contact time between the resin and solution for the
replacing cations to diffuse into the resin and to desorb barium is the
first requirement of the ion exchange desulfating.  However, enough
barium must be in solution to cause the solid to form mainly in the
bulk solution and not in the resin.  This is accomplished by always
maintaining conditions which are favorable for the barium elution
reaction (V) to shift toward the right.  Thus, non-equilibrium conditions
must be maintained.  This environment is created in two ways:
    Other than sulfate

                                -12-

-------
                           TABLE II
                 NORTON FEED WATER PROFILE (*)
                 FROM STATISTICAL TABULATIONS
MINIMUM
Item
Iron (III)
Aluminum
Calcium
Magnesium
Sulfate
PPM
26
11
37
14
270
N x 10J
2.60
1.20
1.85
1.15
5.60
MEAN
PPM
113
35
104
37
904
N x 10
10.0
3.4
5.2
3.0
18.8
MAXIMUM
PPM
229
72
152
88
1650
N x 10-3
20.3
8.0
7.6
7.2
34.4
(*)  A combination of Grassy run and a side tributary.
                               -13-

-------
1.  Split the resin feed among the exchangers so that freshly
    loaded resin is being constantly introduced over the entire
    system.  This will increase the amount of barium eluted
    throughout since a high fraction of barium on the resin is
    maintained over the entire system.

2.  Introduce feed solution to all the  exchangers at a pre-set
    ratio.  This monitors the ratio of  "other" anion/sulfate in
    order to perpetuate the barium sulfate reaction (VI) in
    solution throughout the system.  Therefore, equilibrium
    between the barium on the resin and barium in solution is
    never attained due to the constant  need for the latter because
    of a continuous introduction of sulfate bearing solution.

Elution Rate Studies

The rate of elution of barium from Amberlite(*) IR-120, 14 to 50 mesh,
120F, 40 to 50 mesh, both 8% cross-linked, and IR-118, 14 to 50 mesh
particle assay, 4% cross-linked resin was studied in small scale tests
using brine solutions of varying normality.  The purpose was to study
the various parameters of the equilibrium part of the model and to
determine the degree of elution with normalities as low as 0.003N Na  ,
and 0.002N Ca  , which are approximately the conditions of lean Norton
feedwater (see Table II).  The tests were conducted by vigorously
mixing 42 ml. of wet settled barium loaded resin with one liter of
solution for 20 minutes.  The slurry was separated by filtration; then
the same resin was mixed with a fresh portion of solution for the same
period as before.  The test period was  generally 120 minutes (6 cycles).

After initial problems due to improper  resin loadings were overcome,  the
following were affirmed as is evidenced in Table III:

1.  The resin is highly selective toward calcium over sodium; the
    selectivity range is 2.0 - 3.9(**); this increases with the
    elution time and the decrease of barium on the resin.

2.  There is a slight advantage of the  higher available exchange
    area of the more open 4% cross-linked resin over the 87 cross-
    linked.  The specific loading of the coarse, open resin was
    measured as 180-185 gram barium per liter of wet barium loaded
    resin compared to 215-225 for the fine, 8% resin.

3.  Earlier tests which were run during adverse loading conditions
    indicated that the fine particle assay resin (40 to 50 mesh)
    had a slight advantage over the coarse (14 to 50 mesh) resin.


(*)   Rohm & Haas Co. trademarks for strong acid cation exchange resins.
      Mention of commercial products does not imply endorsement by the
      FWQA.

(**)  Sodium taken as the base value (1.0).


                               -14-

-------
TABLE III
BATCH
SOLUTION
0.1N CaCl2
0.1N CaCl2
0.1N NaCl
0.1N NaCl
0.05N CaCl2
0.05N CaCl2
0.05N NaCl
0.05N NaCl
0.02N CaCl2
0.02N CaCl2
0.02N NaCl
0.02N NaCl
0.01N CaCl2
0.01N NaCl
0.005N CaCl2
0.005N CaCl2
0.005N NaCl
0.005N NaCl
0.002N CaCl2
0.002N CaClo
ELUTION OF BARIUM BY BRINE SOLUTIONS
RESIN TYPE
IR-118
IR-120 F
IR-118
IR-120 F
IR-118
IR-120 F
IR-118
IR-120 F
IR-118
IR-120 F
IR-118
IR-120 F
IR-118
IR-120 F
IR-118
IR-120 F
IR-118
IR-120 F
IR-118
IR-120 F
PERCENT
One Cycle
27
22
10
8
18
15
6
5
11
9
3.2
2.7
7.3
2.0
4.2
3.3
1.4
1.1
2.6
1.8
BARIUM ELUTED
Six Cycles
72
58
25
23
58
47
16
13
41
33
8.6
7.2
27.4
5.8
19
15.5
4.6
3.3
10.4
7.8
        -15-

-------
Figures 12 to 15 in Appendix C contain more detailed information on
the elution of barium from sodium and calcium chlorides using IR-120F
and IR-118.  In these figures, the fractions of the initial barium
concentration remaining in the resin, are shown as a function of the
number of cycles (20 minutes per cycle).

These results can be applied to a system in which a sulfate bearing
brine is used.  A 0.02N NaCl solution could elute less than 9% of the
barium in 120 minutes, and 0.005N NaCl could elute only half that
amount.  If a 0.005 sulfate normality were present with the 0.005N
chlorides (1:1 ratio), enough barium would be eluted from the resin in
120 minutes for complete desulfation.  This, however, is assuming that
a large excess of resin is used, with fresh feed solution to six
separate equilibrium stages (cycles).  However, for a 0.02N NaCl solution,
only about 0.01N barium could be eluted in 120 minutes, which would
result in only 50% desulfating in the above systems.  Therefore, it
was thought that in addition to large excesses of resin, it would be
necessary to introduce fresh resin feed at certain stages of the cycle.
Using the same theoretical equilibrium model, fresh resin feed to the
first five stages would be sufficient for 1007 desulfating for a 0.02N
sulfate solution.  However, as the chloride/sulfate ratio decreases
below 1, complete desulfation becomes impossible in the range under
consideration.  This range coincides with the mean and lean Norton
normalities.

Also included in Appendix C are Figures 17 to 20 which demonstrate the
absolute amounts of barium eluted in each successive equilibrium stage.
The slope of these curves between successive stages (cycles) gives an
indication of the relative feed solution splits which would be used in
a continuous operation having a staged cascaded ion exchange arrange-
ment with split stream feeds.   However, due to use of split resin feeds
and varying sulfate concentrations in the actual system, the optimum
proportion must be determined experimentally.

Figure 21, Appendix C, shows the effect of using bicarbonate in place
of chloride anions in the elution of barium.  The use of bicarbonate
ions therefore appears to be comparable, though slightly less favorable
due to decreased mobility.

Figures 2 to 5 show the equilibrium curves for these runs.   They exhibit
the increase in favorable conditions with increasing normality-  and
with calcium over sodium ions.  The upper dotted line represents the
fully loaded barium resin.  It is noted that as the concentrations
decrease, the curves become more hyperbolic.  This allows a clearly
defined cut point to be determined.  At this point, the maximum
economical theoretical stages  is established.  Beyond this  point, each
                               -16-

-------
   240
             0.005 N
   200 -
a)
4J
-I
1-1

-------
200
               0.2         0.4         0.6         0.8

                   Barium in Solution, grams per liter

      Figure 3:    Equilibrium Between Barium Resin and
                   Solutions of Sodium Chloride (IR-118)
1.0
                                          CCCo. Job 41520
                                          U. S. Dept. of Interior
                                          FWQA Cont. 14-12-518
                          -18-

-------
 -0.002 N  [-0.005 N
             Barium in Solution, grains per liter

Figure 4:  Equilibrium Between Barium Resin and
           Solutions of Calcium Chloride (IR-120F)
                                       CCCo. Job 41520
                                       U. S. Dept. of Interior
                                       FWQA Cont. 14-12-518
                         -19-

-------
               Barium in Solution, grams per liter
Figure 5:
Equilibrium between Barium Resin and
Solutions of Calcium Chloride (IR-118)
                                      CCCo. Job 41520
                                      U. S. Dept. of Interior
                                      FWQA Cont. 14-12-518
                         -20-

-------
successive equilibrium stage results in only a small additional amount
of barium in solution.  Thus, an increasingly large number of stages
is required to bring about relatively small increases in desulfating
efficiency.  These curves again illustrate the desirability of using
a small number of stages with a large excess of resin in each stage.

Continuous Tests

In the continuous tests, the desulfating of feeds consisting of sulfate
bearing brines was attempted.  Thus, the complete desulfating model was
studied together.

The apparatus used resembles the one described in the Bureau of Mines
Report (2).  It is shown in Figure 16, Appendix C.  The main difference
is in the addition of the split resin feed.  The variables studied were
resin type, normality of the mother' liquor, resin and feed splits, and
the molar ratios of both BaiSO^ and C
In the operation, a loaded, washed resin was continuously fed by a
peristaltic pump to the first bucket (*) at 2 to 4cc/min.  The synthetic
sulfate bearing brines were continuously metered to each of the tanks
through individual rotameters at a combined rate of 170 cc/min. to provide
an overall residence time of 110 to 120 minutes.  The plastic tanks were
equipped with variable speed agitators and were covered to avoid resin
loss due to splashing.

At startup, the first tank was partially filled with a small amount
(1000 cc) of a dilute sodium or calcium chloride solution which was
usually of the same normality as in the feed solution so as not to
mask the results.  A small amount (6 to 24 cc) of barium loaded resin
was then mixed for 20 minutes with this solution in order to initiate
the reaction.  Thus, by eluting barium from the resin, an initial con-
centration of barium in the feed solution was assured.

The feed solution flow was then started to the first tank.  After the
overflow from the first tank filled the second, the brine feed to the
second was turned on, and so on down the line.  Samples were taken from
each exchanger at one hour intervals from the time of overflow of that
tank.  The samples were tested for pH and sulfate.  The resin feed rate
was measured regularly, but was found to vary greatly during short periods
of time due to the cyclic pump characteristics, and the average values
were used.  However, the solution feed rates were kept constant.


(*) A second pump was later installed in order to feed resin to
    exchangers 2 and 3.
                               -21-

-------
Table IV is divided into two parts; the "A" section summarizes those
runs in which the ratio of the sulfate to chloride normality is 1.
The "B" section is composed of those runs in which the ratio is greater
than 1.  In general, the table shows that equilibrium conditions are
favorable for those solutions which are of type "A."  Furthermore, 4 x
the stoichiometric amount of resin was a definite improvement over 2 x
the amount, for the runs in which the -40 to +50 mesh 8% cross-linked
resin was used.  In "A" type runs where calcium salt was used, a Ba:S04
ratio of 2 was sufficient; however, when sodium ion was used in place
of some of the calcium, a ratio of 4 was required for 100% desulfating.

For the cases where the ratio of 804:0! was greater than 1, designated
as "B" solutions, the desulfating percent dropped notably as the ratio
was increased.  The use of a higher resin rate did not have the signif-
icant effect that was noticed for the type "A" solutions; however,
splitting the resin feed to two tanks improved the efficiency threefold
when the feed solution splits were also adjusted.  It was believed at
this time that proper adjustment of this split in combination with a
resin feed to three tanks, using a high ratio Ba:S04 could result in
high desulfating efficiencies for solutions in which the 804:Cl ratio
was not so favorable.  This could significantly reduce the need for
pressurization or bicarbonate addition.  Further experimentation would
be needed to verify these theories, in either pilot or bench scale
operations.

Filtration

A settling test, using Neptune microfloc inclined tubes instead of a
clarifier reaction zone, indicated that the settling properties of
neutralized AMD would be improved.  A 2.5 liter synthetic Norton type
solution of 117 ppm Fe+ , 35 ppm Al+ , and 100 ppm Ca+^, and an iron
rich solution of 1000 ppm Fe+3 were both neutralized by soda ash.  A
3-inch by 3-inch square tube proved to be adequate for the lean solution
and a five fold increase in settling rate over "quiet" settling was the
result.  However, the iron rich sludge filled the vertical section at a
pumping rate of 50 ml/min.  This was due to inherent limitations of the
"as built test unit," and the run was cancelled.  However, at a relative-
ly small additional cost in comparison to that of the clarifier, the
tubes could be installed in the reaction zone of a conventional clarifier,
and this would insure a high clarity overflow liquor.
                               -22-

-------
                                                  TABLE IV
                                  CONTINUOUS ION EXCHANGE DESULFATING RUNS

                                    A:   Solutions with SO^/Cl"  =  1.0
TEST
                    819
822
          826
829
                                                            912
918
919
                    922
                                                   926
Solution Normality,
Sodium Sulfate
Sodium Chloride
Calcium Chloride
Total Solution
Feed, cc/min.
Stage 1
Stage 2
Stage 3
Stage 4
Resin Type
Load ing, gm/ liter
Total Resin
Feed, cc/min.
Stage 1
Stage 2
Equivalents -
BA:S04
Equivalents -
504: ci
Run Duration, min.
Sulfate
Concentration , ppm
Stage 1
Stage 2
Stage 3
Stage 4
Initiant - lOOOcc
Norraality,CaCl2
Resin, cc
0.01
-
0.01

170
90
50
20
10
IR-120F
225

2
2
-

3.86

1
240


99.6
99.5
72.2
92.1

0.05
18
0.01
-
0.01

170
90
50
20
10
IR-120F
225

4
4
-

7.73

1
210


93.5
92.3
93.8
93.5

0.05
18
0.01
-
0.01

170
90
50
20
10
IR-120F
225

2
2
-

3.86

1
395


97.9-
88.0
76.4
68.4

0.05
18
0.01
0.01
_

170
90
50
20
10
IR-120F
225
-
2
2
-

3.86

1
200


100.0
86.6
81.5
85.6

0.10
18
0.005
-
0.005

170
90
50
20
10
IR-118F
181

2
2
-

6.22

1
350


96.9
100.0
100.0
100.0
(**)
^ ' 0.01
11.8
0.005
-
0.005

170
90
50
20
10 (*)
IR-120C
237

2
2
-

8.03

1
360


98.4
98.4
99.0
99.0

0.01
11.8
0.01
0.005
0.005

170
90
50
20-
10
IR-120F
228

2
2
~

3.90

1
240


99.4
99.4
99.4
99.4

0.01
11.8
0.01
0.008
0.002

170
90
50
20
10
IR-120F
228

2
2
~

3.90

1
240


63.3
50.0
47.9
43.8

0.01
11.8
0.01
0.008
0.002

170
90
50
20
10
IR-120F
228

4
4
~

7.80

1
350


100
100
100
100

0.01
24.0
(*)   Rohm & Haas Co. trademark for strong acid cation exchange resin of the gel type with -14 to +50 mesh
      particle assay (commercial grade) and 87, cross-linked.

(**)  NaCl

-------
                                                       TABLE  IV
                                       CONTINUOUS  ION  EXCHANGE  DESITUATING RUNS
.p-
i
TEST
Solution Normality
Sodium Sulfate
Sodium Chloride
Calcium Chloride
Total Solution
Feed, cc/min.
Stage 1
Stage 2
Stage 3
Stage 4
Resin Type
Load ing ,gm/ liter
Total Resin Feed,
cc/min.
Stage 1
Stage 2
Equivalents -
BA:S04
Equivalents -
SO^ :Cl
Run Duration, min.
814

0.01
-
0.005

170
90
50
20
10
IR-118
187

2
2
-

3.08

2
110
B:
820

0.01
-
0.005

170
90
50
20
10
IR-118
181

4
4
-

15.2

2
276
Solutions
903

0.03
-
0.002

170
90
50
20
10
IR-120F
225

6
6
-

3.86

15
335
with SO^
908

0.01
-
0.002

170
90
50
20
10
IR-120F
225

2
2
-

3.86

5
325
/Cl" >
909

0.005
-
0.002

170
90
50
20
10
IR-120F
228

1
1
-

3.88

2.5
335
1.0
924

0.01
-
0.005

170
40
70
40
20
IR-118
187

4
4
-

6.40

2
240
930

0.01
-
0.005

170
40
70
40
20
IR-120F
228

4
2
2

3.9/7+

2
225
1001

0.01
-
0.005

170
90
50
20
10
IR-120F
228

4
2
2
f Jf\ ( 
( > 3.9/7+

2
345
1006

0.01
-
0.005

170
70
70
20
10
IR-120F
228

4
2
2
^,\
; 3.9/7+

2
240
Sulfate Concentration,
ppm
Stage 1
Stage 2
Stage 3
Stage 4
Initiant - lOOOcc
Normality, CaCl2
Resin, cc

24.3
12.8
17.9
21.4

0.01
11.7

23.0
17.2
14.8
17.2

0.05
18

8.76
-
-
5.88

0.03
35.4

20.1
-
-
12.3

0.01
11.8

32.9
15.7
28.6
41.0

0.005
5.9

100.0
48.0
26.5
26.0

0.01
23.6

99.4
100
67.8
64.6

0.01
24.0

19.8
19.8
20.8
19.8

0.01
24.0

37.5
33.3
35.4
29.2

0.005
24.0
                                                                                                                     (*)
         To Stage 1; Additional Resin added  to Stage  2 i;l  3.9  BA:SO^ to bring total to over 7.

-------
                          ACKNOWLEDGMENT
This work was completed by Catalytic, Inc. of Philadelphia in
accordance with the FWQA Contract.  All efforts were under the
direction of Mr. John Gilmore, Project Manager.  Process develop-
ment, preliminary design and economic evaluation, laboratory
guidance and progress and final report preparation were completed
by Gideon P. Gelblum, lead process engineer, with assistance from
William H. Weber, process engineer and other Catalytic personnel.

The experimental evidence, which is a part of this report, was a
direct result of the investigations which were conducted at the
laboratories of Cyrus Wm. Rice and Company of Pittsburgh, Penn-
sylvania under the supervision of Mr. William E. Bell, Manager of
Operations Research.

We also wish to acknowledge the cooperation of the Neptune Micro
Floe, Inc. of Corvallis, Oregon in the use of their pilot tube
settler equipment.

We wish further to thank the many companies, whose quotations and
other valuable information contributed to the designs and economic
evaluations.
The objective of this project was to investigate a potentially low
cost  treatment method for mine drainage pollution control.  Although
the results of this investigation indicated the treatment method
could possibly be successful, the high cost of purified water as
indicated by the first Phast study dictated that the method not be
investigated further.  This project of EPA was conducted under the
direction of the Pollution Control Analysis Section, Ernst P.. Hall,
Chief, Donald J. O'Bryan, Jr., Program Manager, and Robert B. Scott,
Project Officer.
                               -25-

-------
List of manufacturers contributing to




         Name




Worthington Corporation




Neptune Meter Co.




Graver Water Conditioning Co.




Sweco, Inc.




Eastern Industries




Vanton Pump & Equipment Corp.




Dorr-Oliver, Inc.




Pennwalt Corp.




Cleveland Mixer Co.




Mixing Equipment Co., Inc.




AMF Cuno Division




FMC, Corp.




Mitchell Plastics, Inc.




Agile Division: Walge Co.




International Salt Co.




Penna. Pump & Compressor Co.




Bartlett-Snow Division




Rex Chainbelt, Inc.




Pulverizing Machinery Division




Meyer Machine Co.




Rohm & Haas Co.




Croll-Reynolds Engineering Co., Inc.
the equipment cost estimate:




         City,  State




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Homer City,  Pa.




      Pittsburgh,  Pa.




      Philadelphia, Pa.




      Clarks Summitt,  Pa.




      Easton,  Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      Philadelphia, Pa.




      San Antonio, Texas




      Philadelphia, Pa.




      Stanford,  Conn.
                               -26-

-------
                            REFERENCES
1.  "Mine Drainage Treatment - State of the  Art  and Research  Needs,"
    R. Hill, U. S. Dept.  of Interior, FWPCA, Mine  Drainage  Control
    Activities, Cincinnati, Ohio,  Dec. 1968.

2.  "Preliminary Process  Development Studies for Desulfating  Great
    Salt Lake Brines and  Sea Water," D. R. George,  J. M. Riley  and
    L. Crocker, Bureau of Mines RI 6928,  U.  S. Department of  the
    Interior, 1967.

3.  U. S. Dept. of Interior; Office of Saline Water,  Spec.  1822,
    CCCo. Contract 41070, April 26, 1968, O.S.W. Contract
    14-01-0001-1175 (75,000 GPD Desulfating  Pilot  Plant).

4.  U. S. Dept. of Interior, O.S.W., R & D Prog. Report No. 289,
    "Evaluation of Brine  Desulfating Process as  Applied to  Desali-
    nation Phase 1, Oct.  1967-"

5.  "Conception, Design and Installation of  a System for Treatment
    of Acid Mine Drainage," U. S.  Dept. of Interior,  FWPCA  RFP  No. WA
    68-144, Catalytic, Inc., June  7, 1968.

6.  Progress Report #2, CCCo. 41520, FWPCA Contract 14-12-518,
    September 10, 1969.

7.  Letter, Scott, R.  B., FWPCA, to Gilmore, J.  F., Jr., Catalytic,
    Inc., June 16, 1969.

8.  Browning, J. E., "Freshening Acid Mine-Waters," Chem. Eng.,
    Jan. 12, 1970, p.  42.
                               -27-

-------
                            APPENDICES
A.  Statistical Analysis of AMD

    Table V    -  Norton Site - Composite
    Table VI   -  Norton Site - Stream G-l-A
    Table VII  -  Norton Site - Stream G-l
    Table VIII -  Hollywood Site - Proctor 1
    Table IX   -  Hollywood Site
    Table X    -  Hollywood Site
    Table XI   -  Hollywood Site
    Table XII  -  Hollywood Site
    Table XIII -  Hollywood Site
    Table XIV  -  Hollywood Site
B.
                               Proctor 2
                               Tyler Run
                               Bennett's Branch
                               Stream Combination No.  1
                               Stream Combination No.  2
                               Composite
Pretreatment Tests
    Table XV
    Table XVI  -

    Figure 6

    Figure 7

    Figure 8

    Figure 9

    Figure 10  -
    Figure 11  -
              Summary of Pretreatment Tests
              Sludge Settling Tests

              Iron and Aluminum in Solution  after
              Neutralization
              Bicarbonate and pH as  Functions  of Soda
              Ash Addition
              Typical Pretreatment Run for Iron Rich
              Solutions
              AMD Pretreatment:  pH as a Function of
              Mixing Time
              Settling Rates of Synthetic AMD  Solutions
              Settling Rates of Synthetic AMD  Solutions
C.  Batch Elution Tests

    Figure 12  -  Effect of Calcium Chloride Concentrations
                  on Elution of Barium for IR-120F
    Figure 13  -  Effect of Sodium Chloride Concentrations
                  on Elution of Barium for IR-120F
    Figure 14  -  Effect of Calcium Chloride Concentrations
                  on Elution of Barium for IR-118
    Figure 15  -  Effect of Sodium Chloride Concentrations
                  on Elution of Barium for IR-118
    Figure 16  -  Continuous Ion Exchange Desulfating
                  Apparatus -
    Figure 17  -  Elution of Barium from IR-120F Vs Cycles of
                  Elution for Solutions of Sodium Chloride
32

33
34
35
36
37
38
39
40
41
42

43

44
47
48
49

50

51

52
53

54

55

56

57

58

59

60
                               -29-

-------
                            APPENDICES
    Figure 18  -  Elution of Barium from IR-120F Vs Cycles      61
                  of Elution for Solutions of Calcium Chloride
    Figure 19  -  Elution of Barium from IR-118 Vs Cycles of    62
                  Elution for Solutions of Sodium Chloride
    Figure 20  -  Elution of Barium from IR-118 Vs Cycles of    63
                  Elution for Solutions of Calcium Chloride
    Figure 21  -  Effect of Anion on Elution of Barium from     64
                  IR-Resins

D.  One Million Gallon Per Day Plant                            65

    Process                                                     66
       Design Conditions
          Scope
          Basis
       Process Description - Neutradesulfating AMD Process
          Introduction
          Removal of Iron and Aluminum from AMD
          Removal of Sulfates and Carbonates from AMD
          Resin Regeneration
          Effluent Treating
          Waste Treatment
          Barium Sulfate Conversion
          Carbonation
          Sulfur Recovery
       Recovery of Salable By-Products & Logistics
          By-Product Recovery
          Logistics of Neutradesulfating Process
    Plant Equipment                                             80
       Furnace
       Reactors
       Towers
       Heat Exchangers
       Vessels
       Pumps
       Materials Handling
       Filters, Screens, Centrifuges,  and other Separators
       Agitators
       Packaged Units
    Plant Layout                                                84
    Economics                                                   85
       Capital Investment
       Operating Costs
       By-Product Sale
                               -30-

-------
                      APPENDICES
Economics (Continued)
   Cost of Neutradesulfated Water
   Expansion of Cost of Neutradesulfating for Larger
     Capacity Plants
   Economic Feasibility of the Neutradesulfating Plant

Table XVII  - Production Plant Daily Traffic Schedule       78
Table XVIII - Production Plant Utilities Summary            79
Table XIX   - Capital Investment for Neutradesulfating      86
              Plants
Table XX    - Economic Evaluation for a 1MM GPD             87
              Neutradesulfating Plant

Figure 22   - Block Diagram for AMD Neutradesulfating       69
Figure 23   - Effect of Plant Capacity on the Unit Cost     92
              of Neutradesulfating
Figure 24   - Effect of Capital Investment on the Unit      94
              Cost of Neutradesulfating

Engineering Diagrams
   Drawing R-211                                            95
   Drawing R-212                                            99
   Layout Sketch                                           102
                        -31-

-------
     APPENDIX A




STATISTICAL ANALYSIS




         OF




         AMD
        -32-

-------
                          TABLE  V
                  STATISTICAL ANALYSIS  OF  AMD
Norton Site
No.
1.

2.
3.

4.
5.
6.
7.

8.
9.
10.
11.
12.
Item
Flow CFS
GPM
MMGPD
PH
ACIDITY
as ppm CaCOo
Fe total, ppm
+2
Fe , ppm
-2
S04 , ppm
HARDNESS
as ppm CaCOj
+2
Ca , as ppm CaCO-j
Ca+2/ HARDNESS
M + , as ppm CaC03
, . +3
Al , ppm
CONDUCTIVITY, Micro-
Mean
6.64
2,980.
4.29
2.77
609.

113.
4.13
904.
411.

259.
0.68
152.
35.3
1,570.00
- Composite
Standard
Deviation
6.25
2,800.
4.04
0.17
180.

39.7
1.42
263.
120.

72.3
0.54
47.6
10.7
A27.
Popu-
lation
91
-
82
82

82
22
82
82

40
40
-
40
82
Minimum
0.80
359.
0.52
2.50
83.1

26.0
2.00
270.
151.

92.0
0.59
59.0
11.0
150.
Maximum
23.0
10,300.
14.9
3.30
947.

229.
7.20
1,650.
744-.

380.
0.86
364.
72.0
2,250.
mhos "J 25C
                         -33-

-------
                                TABLE VI
                      STATISTICAL ANALYSIS OF AMD
                       Norton  Site -  Stream G-1A
No.
1.
2.
3.
4.
5.
6.
7.
Item
Flow CFS
GPM
MMGPD
PH
ACIDITY
as ppm CaC03
Fe total , ppm
Fe , ppm
SO^" , ppm
HARDNESS
                           Mean
                                 Standard   Popu-
                                 Deviation  lation  Minimum  Maximum
 9.

10.

11.

12.
  as ppm CaCO-,

Ca  ,  as ppm CaCO-j

Ca+2/HARDNESS
Mg+ ,as ppm CaCO

  +3
Al  , ppm
                -,
     CONDUCTIVITY,  Micro-
      mhos  O 25C
    7.94      15.2
3,560.      6,800.
    5.13       9.80

    2.77

  525.5


  115.



  717.

  340.


  231.

    0.70

  109.

   31.0

1,280.
                                          0.13

                                        216.


                                         44.7



                                        211.

                                         98.20
                                              26



                                              26

                                              26


                                              26



                                              26

                                              26
                                                           0.80       80.
                                                         359.      35,900.
                                                           0.52       51.7
                                                      2.60

                                                     83.1


                                                     26.0



                                                    270.

                                                    151.
 69.30     26     92.0

  0.10     26      0.60

 28.9   #7 - #8   59.0

  9.10     26     11.0

482.       26    150.
    3.10

  947.


  210.



1,120.

  492.


  330.

    0.90

  162.

   45.

1,900.
                               -34-

-------
                          TABLE VII
STATISTICAL ANALYSIS OF AMD
Norton Site 
No.
1.

2.
3.

4.
5.
6.
7.

8.
9.
10.
11.
12.
Item
Flow CFS
GPM
MMGPD
PH
ACIDITY
as ppm CaC03
Fe total , ppm
+2
Fe , ppm
S04" 2, ppm
HARDNESS
as ppm CaCO.j
+2
Ca , as ppm CaCC>3
Ca+2 /HARDNESS
Mg+2 , as ppm CaC03
Al+3, ppm
CONDUCTIVITY, Micro-
Mean
5.80
2,590.
3.74
2.80
632.

105.
4.13
990.
449.

312.
0.7C
137.
41.5
1,700.
 Stream G-l
Standard
Deviation
6.31
2,830.
4.10
0.20
141.

26.5
1.43
246.
114.

42.5
0.10
71.1
6.40
326.
Popu-
lation
50
-
50
50

50
22
50
50

14
14
By Dif-
ference
13
50
Minimum
0.76
341.
0.50
2.50
204.

35.0
2.00
410.
167.

243.
0.60
29.
918.
Maximum
23.0
10,300.
14.9
3.30
890.

160.
7.20
1,650.
744.

380.
0.80
364.
51.0
2,250.
mhos @ 25C
                         -35-

-------
                                                      TABLE VIII
i
u>
No.     Item


1.   pH


2.   Acidity



3.   Iron as Fe
                     +2
      4.   Aluminum  (Al+3)


      5.  Sulfate  (SC>4=)


      6.  Difference as Ca


      7.  Conductivity,

          Micromhos @ 20C
STATISTICAL ANALYSIS OF
AMD




Hollywood Site
Proctor 1
Standard
Mean
Equiv-
alents
-
.0028
.0096
.0167
.0043
_
ppm
3.42
462.
77.7
86.2
802.
86.0
1,550.
Deviation
Equiv-
alents
-
-
.0009
.0040
.0054
.0005
.
ppm
0.20
150.
24.7
36.0
260.
10.0
44.6
Popu-
lation
15
15
15
15
15
-
5
Minimum
Equiv-
alents
-
-
.0011
.0014
.0052
.0027
-
ppm
3.72
142.
30.0
13.0
251.
54.0
1,520.
Maximum
Equiv-
alents
-
-
.0043
.0170
.0247
-
-
ppm
3.13
671.
119.
153.
1,190.
-
1,620.

-------
                                                 TABLE IX
No.    Item

1.  pH

2.  Acidity

3.  Iron as Fe
+2
               ,+3,
4.  Aluminum (Al  )

5.  Sulfate (S04=)

6.  Difference as Ca

7.  Conductivity,
    Micromhos @ 20C
STATISTICAL ANALYSIS OF AMD
Hollywood Site
Proctor 2
Standard
Mean
Equiv-
alents
-
-
.0305
.0279
.0724
.0140
_
ppm
3.25
2,640.
853.
251.
3,480.
2.81
3,590.
Deviation
Equiv-
alents
-
.0107
.0117
.0189
.0035
-
ppm
0.10
946.
298.
105.
910.
70.0
2,310.
Popu-
lation
15
15
15
15
15
-
5
Minimum
Equiv-
alents ppm
3.41
1,280.
.0114 317.
.0185 166.
.0466 2,240.
-
3,190.
Maximum
Equiv-
alents
-
-
.0631
.0689
.1337
-
_
ppm
3
5,740
1,760
620
6,420
-
3,770
.10


.00


.

-------
                                              TABLE  X
STATISTICAL ANALYSIS OF
AMD



Hollywood Site
T_yler Run, Pa.
Standard
Mean
No.
1.
i
w 2
00 *'
3.
4.
5.
6.
7.
Item
pH
Acidity
+2
Iron as Fe
Aluminum (Al )
Sulfate (804 )
1 i
Difference as Ca
Conductivity,
Equiv-
alents
-
-
.00005
.0040
.0064
.00235
-
ppm
3.50
149.0
1.50
35.6
309.
47.0
922.
Deviation
Equiv-
alents
-
-
.00003
.0026
.0032
.00057
-
ppm
0.17
75.0
0.90
23.5
156.
11.0
146.
Popu-
lation
15
15
15
15
15
-
5
Minimum
Equiv-
alents
-
-
.00002
.001
.0027
.0017
_
ppm
3.75
65.0
0.50
9.00
132.
34.0
679.
Maximum
Equiv-
alents ppm
3.20
287.
.0001 3.0
.0098 88.0
.0123 590.
-
1,040.
Micromhos @ 20 C

-------
                                             TABLE XI
STATISTICAL ANALYSIS OF AMD
Hollywood Site
Bennett's Branch, Pa.
S tandard
Mean

10
\o
i




No.
1.
2.
3.
4.
5.
6.
7.
Item
PH
Acidity
+2
Iron as Fe
Aluminum (Al )
Sulfate (S04=)
 i
Difference as Ca
Conductivity,
Equiv-
alents
-
.0011
.0034
.0062
.0017
-
PPM
3.80
179.
30.3
30.7
296.
34.0
912.
Deviation
Equiv-
alents
-
.0006
.0025
.0048
.0017
-
ppm
0.50
143.
17.6
22.6
230.
34.0
480.
Popu-
lation
15
15
15
15
15
-
5
Minimum
Equiv-
alents
-
.0003
.0006
.0013
.0005
-
PPtn
4.73
31.0
7.00
5.00
61.0
9.00
376.
Maximum
Equiv-
alents
-
.0021
.0097
.0156
.0038
-
ppm
3.03
462.
59.0
87.0
747.0
76.0
1,340.
Micromhos @ 20C

-------
                                                       TABLE XII
o
STATISTICAL ANALYSIS OF AMD
Hollywood Site
Stream Combination No. 1
Standard
Mean
No.
1.
2.
3.
I
4.
5.
6.
7.
Item
PH
Acidity
+2
Iron as Fe
+3
Aluminum (Al )
Sulfate (S04~)
I i
Difference as Ca
Conductivity,
Equiv-
alents
-
-
.0025
.0104
.0178
.0049
-
ppm
3.34
459.
71.0
94.0
857.0
98.0
1,350.
Deviation
Equiv-
alents
-
-
.0009
.0024
.0039
.0006
-
ppm
0.20
99.0
25.0
22.0
187.
12.0
243.
Popu-
lation
6
6
6
6
6
-
5
Minimum
Equiv-
alents
-
.0012
.0070
.0141
-
-
ppm
3.50
359.
33.0
63.0
679.
-
1,040.
Maximum
Equiv-
alents
-
-
.0034
.0131
.0223
-
_
ppm
3.20
572.
96.0
118.
1,070
-
1,580.
          Micromhos  @ 20C

-------
                                                      TABLE XIII
i
F-
STATISTICAL ANALYSIS
OF AMD




Hollywood Site
Stream Combination No. 2
S tandard
Mean
No.
1.
2.
3.
4.
5.
6.
7.
Item
PH
Acidity
Iron as Fe
+3
Aluminum (Al )
Sulfate (S04=)
Difference as Ca"1"1"
Conductivity,
Equiv-
alents
-
-
.0014
.0062
.0114
.0038
-
ppm
3.35
302.
38.0
56.0
549.
76.0
1,020.
Deviation
Equiv-
alents
-
-
.0007
.0036
.0065
.0022
.
PPtn
0.3
165.
19.0
32.0
310.
44.0
438.
Popu-
lation
6
6
6
6
6
-
5
Minimum
Equiv-
alents
-
-
.0006
.0039
.0051
.0006
-
Ppm
3.80
120.
16.0
35.0
245.
12.0
570.
Maximum
Equiv-
alents
-
.0026
.0128
.0210
-
-
ppm
3.1
543.
72.0
115.
1,010.
-
1,550.
         Micromhos @  20C

-------
                                             TABLE XIV
STATISTICAL ANALYSIS OF AMD
Hollywood Site
(*)
Composite
Standard
Mean

i
.t-
ro
i





No.
1.
2.
3.
4.
5.
6.
7.
Item
PH
Acidity
Iron as Fe+2
+3
Aluminum (Al )
Sulfate (804")
1 i
Difference as Ca
Conductivity,
Equiv-
alents
_
-
.013
.011
.025
.002
-
ppm
3.5
954.
354.
101.
1,220.
30.
1,660.
Deviation
Equiv-
alents
-
.026
.012
.029
-
-
ppm
0.34
1,450.
720.
106.
1,410.
-
1,040.
Popu-
lation
63
63
63
62
62
-
25
Minimum
Equiv-
alents

-
.00002
.0006
.001
.0007
-
ppm
4.74
31.0
0.50
5.00
61.0
14.0
376.0
Maximum
Equiv-
alents
_
-
.166
.069
.134
-
-
ppm
2.90
8,110.
4,620.
620.
6,420.
-
3,770.
Microtnhos @ 20C
(*) Data includes 1 point (6/10 data) not  included  in  separate stream analyses;  this point  is  an
    extreme case and has caused the high standard deviations occurring  in  this composite.

-------
    APPENDIX B




PRETREATMENT TESTS
      -43-

-------
                                                 TABLE  XV
                                       Summary of Pretreatment  Tests
Run Number

Initial pH
Initial Iron, ppm
Initial Aluminum, ppm
Initial Calcium, ppm
Soda Ash Added, gm
Immediate pH
Stoichiometric Amount, gm
Actual/Stoichiometric
Mixing Time, hrs.
  End pH
Sparging Time, hrs.
  End pH
Settling Time, hrs.
  End pH

Analytical Results

pH
Bicarbonate, ppm
Total Iron,  ppm  @ Top
Total Iron,  ppm  @ Bottom
Dissolved Iron, ppm
Total Aluminum, ppm
Dissolved Aluminum, ppm
Total Calcium, ppm
Dissolved Calcium, ppm
Sludge Volume, cc
610B
610D
613A
                                                        613B
                                     613C
                                                                            613D
                                                         616A
                                                                                              616B
                                                                           616C
2.0
1156.
85.
280.
1.752
5.0
1.85
0.95
0.5
-
-
2.
"
5.8
55.
11.2
14.5
9.0
1.14
-
260.
-
150.
2.0
1156.
85.
280.
1.684
5.0
1.85
0.91
0.5
-
0.5
2.

4.7
33.
41.9
45.4
39.5
1.04
-
280.
-
155,
2.1
1156.
85.
-
1.733
5.0
1.85
0.94
2.
5.3
-
2.
5.4
5.1
20.
14.0
18.3
11.7
1.57
0.88
-
-
177.
2.1
1156.
85.
-
2.240
6.1
1.85
1.21
2.
7.3
-
2.
7.6
8.0
642.
1.6
1.9
0.37
0.64
0.20
-
-
180.
2.1
1156.
85.
-
1.773
5.0
1.85
0.96
2.
5.6
0.5
2.
6.5
6.6
20.
1.1
1.6
0.25
0.78
0.15
-
-
165.
2.1
1156.
85.
-
1.910
5.6
1.85
1.05
2.
6.4
0.5
2.
7.0
7.35
80.
2.4
2.5
0.13
1.47
0.05
-
-
125.
2.1
1156.
85.
280.
1.78
-
1.85
1.09
2.
-
0.5
2.
4.85
4.4
57.
43.6
44.8
41.1
2.59
2.45
284.
280.
167.
2.1
1156.
85.
280.
1.88
-
1.85
1.15
2.
-
-
2.
6.40
6.5
97.
2.13
6.71
1.14
1.01
0.92
272.
264.
165.
2.1
1156.
85.
280.
1.99
.
1.85
1.22
2.
.
0.5
2.
7.00
7.0
142.
2.07
3.24
0.38
0.59
0.39
248.
244.
167.

-------
                                                 TABLE XV
                                       Summary of Pretreatment Tests (Cont'd)
Run Number

Initial pH
Initial Iron, pptn
Initial Aluminum, ppm
Initial Calcium, ppm
Soda Ash Added, gm
Immediate pH
Stoichiometric Amount, gm
Actual/Stoichiometric
Mixing Time, hrs.
  End pH
Sparging Time, hrs.
  End pH
Settling Time, hrs.
  End pH

Analytical Results

PH
Bicarbonate, ppm
Total Iron, ppm @ Top
Total Iron, ppm @ Bottom
Dissolved Iron, ppm
Total Aluminum, ppm
Dissolved Aluminum, ppm
Total Calcium, ppm
Dissolved Calcium, ppm
Sludge Volume, cc
                            616D
616E
616F
                   619A
                             619B
                              619C
                                                 619D
619E
620A
2.1
1156.
85.
280.
2.45
-
1.85
1.50
2.
-
-
-
2.
7.7
2.1
1156.
85.
280.
2.22
6.0
1.85
1.36
2.
7.12
0.5
-
2.
7.9
2.1
1156.
85.
280.
2.58
6.4
1.85
1.58
2.
7.77
-
-
2.
8.3
2.0
1156.
85.
-
1.773
4.95
1.85
0.96
2.
5.57
0.5
5.52
2.
6.3
2.0
1156.
85.
_
1.773
5.35
1.85
0.96
1.
5.43
0.5
5.55
2.
5.8
2.0
1156.
85.
..
1.815
5.33
1.85
0.98
1.
5.72
0.5
6.07
2.
6.6
2.0
1156.
85.
..
1.910
5.40
1.85
1.03
2.
6.78
0.5
7.0
2.
-
2.0
1156.
85.
_
1.910
5.41
1.85
1.03
2.
6.95
0.5
7.3
2.
-
2.43
120.
_
_
0.210
5.38
0.171
0.84
2.
5.39
0.5
5.52
2.
4.88
7.5
801.
1.03
2.09
0.25
0.43
0.14
188.
188.
185.
7.6
393.
1.27
2.17
0.44
0.48
0.24
192.
192.
165.
7.8
1015.
0.70
1.88
0.28
0.42
0.22
140.
136.
182.
6.2
39.
2.18
5.3
0.20
0.79
0.08
-
.
105.
5.9
34.
10.4
17.6
5.20
2.16
0.22
-
_
136.
6.6
97.
1.65
3.91
0.50
0.73
0.05
.
_
145.
7.85
302.
3.76
4.58
0.21
0.64
0.13
_
_
125.
7.85
312.
3.27
5.37
0.41
0.47
0.09
_
_
95.
5.0
74.
11.3
13.2
7.7
_
_
_
_
25.

-------
                                                  TABLE XV
                                        Summary  of  Pretreatment Tests (Cont'd)
 Run Number

 Initial pH
 Initial Iron,  ppm
 Initial Aluminum, ppm
 Initial Calcium,  ppm
 Soda Ash Added,  gm
 Immediate pH
 Stoichiometric Amount,  gm
 Actual/Stoichiometric
 Mixing  Time, hrs.
   End pH
 Sparging Time, hrs.
   End pH
 Settling Time, hrs.
   End pH

 Analytical Results

 PH
 Bicarbonate, ppm
 Total Iron, ppm  @ Top
 Total Iron, ppm  @ Bottom
 Dissolved Iron, ppm
Total Aluminum, ppm
Dissolved Aluminum, ppm
Total Calcium,  ppm
Dissolved Calcium, ppm
Sludge Volume,  cc
                             620B
620C
62 OD
623A
                            623B
                             623C
                                                62 3D
                                      702
                                                         703
2.45
120.
35.
-
0.306
5.38
0.255
1.22
2.
5.50
0.5
5.86
2.
5.36
5.6
79.
8.6
9.1
5.9
0.98
0.68
-
-
25.
2.36
120.
35.
100.
0.305
5.36
0.255
1.22
2.
-
0.5
6.17
2.
5.97
6.2
72.
4.2
4.0
0.79
1.09
0.38
96.
96.
25.
2.37
NORTON
FEED-
WATER
0.283
5.45
-
-
2.
-
0.5
6.5
2.
6.33
6.4
76.
3.4
4.1
0.77
1.50
0.21
136.
136.
25.
2.44
NORTON
FEED-
WATER
0.290
5.68
-
-
0.5
5.89
0.25
6.33
2.
6.43
6.9
130.
-
1.1
0.55
0.80
0.30
132.
132.
35.
2.47
NORTON
FEED-
WATER
0.290
5.67
-
-
0.5
5.94
-
-
2.
5.9
6.45
135.
-
2.7
0.94
1.72
0.48
140.
144.
25.
2.36
117.
-
_
0.247
5.75
0.167
1.48
2.
6.82
0.5
6.95
2.
7.02
7.20
142.
-
4.25
0.25
-
-
-
-
20.
2.35
117.
35.
100.
0.334
5.65
0.250
1.33
2.
6.52
0.5
6.62
2.
6.84
7.1
125.
-
3.57
0.73
0.24
0.38
120.
112.
35.
2.32
NORTON
FEED-
WATER
0.508
7.05
_
_
2.
8.51
0.5
8.52
2.
8.39
8.3
1221.
_
3.07
0.09
1.10
0.41
104.
104.
50.
2.40
NORTON
FEED-
WATER
0.319
5.75
_
_
2.
6.54
0.5
6.45
2.
5.05
6.8
121.
_
3.69
0.48
0.44
.005
105.
112.
25.

-------
                    TABLE XVI
              SLUDGE SETTLING TEST!
Cubic Centimeters Occupied By Sludge At Interface
                  Time,  Minutes
Run
Number
610B
610D
613A
613B
613C
613D
616A
616B
616C
616D
616E
616F
619A
619B
619C
619D
619E
620A
620B
620C
620D
623A
623B
623C
623D
702
703
0
510
510
500
500
500
500
505
510
510
510
505
505
500
495
510
495
495
500
500
500
500
500
500
500
500
500
505
5
495

495
493
488
_
477
502
480
502
497
500
485
445
500
380
335
75
100
100
75
_
140
_
-
90
70
10
455
460
485
486
460
453
427
465
412
492
475
480
428
367
493
320
195
25
75
50
25
85
120
-
100
75
70
20
392
382
431
433
399
286
377
405
345
493
395
465
340
265
405
239
143
-
50
-
25
75
75
25
70
50
50
30
327
284
380
370
342
214
335
345
292
445
331
382
282
215
332
193
125
-
50
-
25
60
65
20
55
50
45
40
280
240
330
315
290
180
292
285
255
340
275
330
230
180
270
170
120
-
-
-
-
50
50
20
50
50
40
50
245
244
288
271
250
165
255
246
230
287
242
280
200
155
235
163
110
-
-
-
-
-
-
-
-
50
35
60
215
210
254
237
225
155
231
220
215
255
225
247
175
147
205
150
107
-
-
-
-
35
25
20
50
50
25
120
150
161
17D
180
165
125
167
165
167
185
165
182
105
136
145
125
95
-
-
-
-
35
25
20
35
50
25

-------
26
                                                               10
                               PH
    Figure 6:  Iron and Aluminum in Solution
               after Neutralization
                                          CCCo. Job 41520
                                          U.S. Dept. of Interior
                                          FWQA Cont. 14-12-518
                           -48-

-------
  200C-
  16CC
0)
4J

-------
o
                 5.0
                 4.0
                 3.0
                 2.0
                 1.0
                                  1.0
          Run  #/09A,  Initial  Concentraticns=
                   t:'e+;T
                   Al
                 400
                      
-------
5.5
5.4
5.3
5.:
5.1
5.0
 4-.
Ratio of Soda Ash added to
the Stoichiometric Quantity.
Initial Concentrations 
          1156ppm
               Fe
               Al
                                 +3
                                        85ppm
           20
      Figure  9:
     40       60
       Time, Minutes
           80
100
120
AMD Pretreatment;
pH as a Function of Mixing Time
                                     CCCo.  Job  41520
                                     U.  S.  Dc.pt.  of Interior
                                     FWQA Cont.  14-12-518
                       -51-

-------
    500
O
U
0)
U

-------
              500
             400
i
in
         O
         u
         0
         u
         IH
         t:
         4J
300
              200
         OT

         W
         O
         o
              100
                         The Initial Concentration is
                         (120 ppm Fe"1"^, 35 ppm
                   Figure  11:
                                          60          80
                                  Settling Time^  Minutes
                 Settling Rates of  Synthetic AMD Solutions
100         120
  CCCo. Job 41520
  U.S. Dept. of Interior
  FWQA Cont. 14-12-518

-------
    APPENDIX C




BATCH ELUTION TESTS
        -54-

-------
ttf
PQ
u
cfl
J-l
          O.b
          0.5 -
          0.4
                                     80

                 Time, Minutes


12:   Effect of Calcium. Chlorj.de Concentrations

     on Elution of Bariun.  for IR-120V
                                                                          LOU
TZCT
                                                                  CCCo.  Job 41520

                                                                  U.  S.  DcjL.  of Interior

                                                                  1VQA Con.t .  14-12-518

-------
en
0)
OS


O
tfl

,I
cd
O


O
O
*
S-l
         0.8
         0.7
         0.6
        0.5
                      -zfr
                                                          -sir
                                       Time.  Minutes
               Figure  13:   Effect  of Sodjui" Chloride Concentrations
                            on  Elution of BarJur, for IR-120?
                                                                 CCCo.  Job 41520

                                                                 U.  S.  Drpt.  of Interior

                                                                      Cont.  In-12-518

-------
01
J
I
         c
         ~t-l
         U3
         0!



         5


         3
         H

         cfl


         r-(
         cfl

         iJ

         5
         o
         cO
         1-1
                 0.5  -
                 o.;
                         Fi.rure  14:
                                                                      80
          Time, Minutes



Effect of Calcium Chloride Concentrations

on Elution of  Bariuir; for IR-118
                                             100
120
                                                                             CCCo.  Job  41520

                                                                             U. S.  Dept.  of Interior

                                                                                  Cont.  14-12-518

-------
I
in
00
          c

          BJ
          o
          n
          d
          ca
          r-l

         4J



         C

         Hi
          O
          o
          c<)
          l-i
1.0



0.9




0.8




0.7





0.6
                  0.5
                                 20
                          40           60

                             Time, Minutes
80
100
                                                                                             120
                      Figure  15:   Effect of Sodium Chloride  Concentrations

                                   on Elation of Barium  for IR-118
                                                                           CCCo. Job  '-^1520

                                                                           U. S. D-^pt.  of InL.'.rior

                                                                            VQA Cont.  1-'.-12-518

-------
  Feed
Solution
                                  Loaded Res in
                                    I   ( + Water)
V-101       Sulfate Feed Tank

V-102       Resin Feed Hopper

V-105 to 8  Desulfating Reactors

P-101       Resin Displacement Punp

            Spent Resin Bucket

            BaSO^ - water slurry bucket

            Vibrating Screen
V-103

V-lOA

H-101

A-101
CCCO. 41520
U.S. Dept.  of Interior
FWQA  Cont. 14-12-518
            Paddle Agitators (A) - two blade

           FIGURE 16- CONTINUOUS ION EXCHANGE DKSULFATING APPARATUS
                                       -59-

-------
30
                           3            4
                    Ci. cles  of  Elution

  "jj-ure  17:  Elution of Barium from
    .R-i20' for Solutions of Sodium
  Chloride
CCCo. Job 41520
U. S. Dept. of Interior
"WQA Cont. 14-12-518
                        -60-

-------
I
o
o
CO
c
H
H
M
    200  -
    100
          Figure  18:
         345

      Cycles of Elution

Elation of Barium from IR-12U?' for
Solutions of Calcium Chloride
                                          CCCo  Job  41520
                                          U.  S. Dept.  of  Interior
                                          FWQA  Conr.  14-12-518
                            -61-

-------
30 ill
2 3 4
C'-oies of Elution
' ure IV  Elut ,011 of Bai i un f^om
.R-i't-S foi Solutions of
Sodiiu" Chlov'de

9^^>
5 6

CCCo. Job .',1520
U. S. Dc-^L. of Interior
t'WQA. Gout. L',- 12- 5 18
-62-

-------
k-<
o
o
to
(t
W
1000
 900

 800

 700

 600

 500


 400



 300
    200
    100
                                 3            4            f

                              Cycles  of Elution

            Figure  20:   Elation  of  Barium from IR-118  for
                        Solutions of  Calcium  Chloride
                                          CCCo.  Job 41520
                                          U.  S.  Dept.  of Interior
                                          FWQA. Cont.  14-12-518
                               -63-

-------
          200
3
O
cd
03
100
 yo
 80
 70
 oO

 50 U

 40
               01            23            4            5
                                          Cvcles  of  Elution

                    Figure 21:   Effect of Anion on Elution of  Bariun  from  IR-Rasins
                                                               CCCo.  Job  41520
                                                               U.  S.  Dept  of Interior
                                                               FWQA Conr.  14-12-518

-------
    APPENDIX D




ONE MILLION GALLON




   PER DAY PLANT
        -65-

-------
                PRODUCTION PLANT - ONE MILLION GPP

                              PROCESS
Design Conditions
The specified production plant is based on experience from several
sources, plus additional laboratory experimentation.  The AMD
pretreatment and desulfating sections are based on experimental
laboratory work (see page 7) conducted under the FWQA contract.
The resin regeneration, effluent treatment, barium conversion, and
carbonation sections are based mainly on the work conducted by the
Bureau of Mines (2)., and on previous Catalytic design packages
(3, 4).

Basis:

The design throughput of the production plant is one million gallons
per day of acid mine drainage.

The unit is designed to handle feed of the composition listed in the
Appendix A, Table V, for Norton Profile, but is readily adaptable' '
to Hollywood Profile (listed in Appendix A, Tables VIII to XIII).
These are statistical compilations from available data.  Material
balance and economics are based on the Norton Site.

The ion exchange regenerant will be an aqueous solution of 50-50
barium hydrosulfide and barium hydroxide, which has a normality of
0.1.

The production plant shall be fully automated so that it can be run
continuously to satisfy commercial requirements.

The ion exchange regeneration process shall be a continuous process,
attaining the maximum allowable superficial velocity through T-301 A/B
Loading Columns and minimizing barium leakage into the effluent solution.

Complete removal of barium from the pore solution of the resin is
required (T-302 Rinsing Column).

Attrition of ion exchange beads shall be kept to an absolute minimum
throughout the production plant.
 (*) Additional clarifier area is required.

-------
The recovery of high clarity AMD from pretreatment and desulfating
operations shall be maximized by use of special flocculents, reaction
type gravity-settling devices, and subsequent dewatering by centrif-
ugation.  Carry over of barite fines from clarification shall be
minimized and filtration shall be used to insure a clear product.

A relatively high excess of resin over the amount needed stoichio-
metrically for complete desulfating will be used due to low
reactivity of the resin in dilute ionic solutions (laboratory results
dictate use of 4x the stoichiometric quantity).  A 20-30 percent ex-
hausted resin must be regenerated, requiring use of a dilute liquor
in order to minimize barium losses and effluent treating costs.

The following criteria were used in compiling a capital cost estimate
for the economic evaluation.

All utilities will be made available on the premises.

Makeup regenerant in the barite form will be kept in dry storage
(50-pound bags).  Sulfuric acid will be stored in 42-gallon barrels.
Sodium chloride (Salt Cake) shall be stored in a lixator vessel which
generates a saturated solution to the process.  Flocculent for both
pretreatment and desulfating will be received and stored in bags.
Makeup resin will be purchased and kept in dry storage containers.
Storage capacity will b for two week production periods for all
chemicals, and for sulfur by-product.

A laboratory and equipment, operating and control room, office,
maintenance, and storage facilities shall be furnished.

Loading and unloading facilities and roadways shall be supplied.

The makeup filter, feed pump, and storage lixator shall be furnished
with concrete bases, and the process water surge pond shall be
fiberglass or polyethylene lined.

The plant, facilities, and offsites require a plot 350 x 200 feet wide,
or 1.61 acres.
                               -67-

-------
Process Description - Neutradesulfating AMD Process

introduction

The process flowsheets (R-211 and R-212, Appendix D) are based
on the neutradesulfating of one million gallons of acid mine
drainage per day (1MM GPD).   The process described here is
shown in block form in Figure 22 and is a modification of the
process developed by the Bureau of Mines (2).  This process
would provide a product which would meet interstate water
quality criteria.

The production plant is a continuous, automated operation
equipped with Pretreatment,  Desulfating, Resin Regeneration,
Effluent and Waste Treatment (drawing R-211), Barium Conversion
and Carbonation Sections (drawing R-212), and By-Product Recovery^
The desulfating section is pressurized.

Removal of Iron & Aluminum from AMD

Raw AMD is passed through a traveling water screen prior to
pumping in order to remove twigs, leaves and other debris from
the feed.  The screen, along with the AMD feed pump, is in a
concrete surge pond located in the feed stream itself.  Filtered
AMD is then distributed in parallel to two pretreatment reactors.
Prior to entering the reactor,  the AMD feed stream is injected
with a dilute sodium bicarbonate solution from the Carbonation
Section at a rate which is controlled by the pH of the pretreated
AMD stream.  After a period (generally one hour) of air sparging
in vigorously agitated pretreatment reactors to insure both a
complete pretreatment reaction and oxidation of all ferrous ion
to ferric, the neutralized mine drainage flows by gravity into the
feedwell of the pretreatment clarifier.  The iron suspension is
then passed into an agitated reaction zone where it is contacted
with an anti-flocculent which aids in decreasing spherical inter-
ferences between the particles  which can cause a colloidal state
(*) Sulfur Recovery Plant is a packaged item and the flow diagram
    and process description are not included here.
                               -68-

-------
OW051165
        NFTTP..IIZATION

       Air          Floe
TON EXCMKCE -  DESL'LFATING


   Flue Gas
                                         C02

                                          uegasit i-
                                                     .JBarluni|  juegasit L
                                                    J^Completion I  1  rMon
^-^   A	     I	,   ,	*     iM,in        _!_J-^H
( AMI) \-|pretreatnient|-p) Sett ling  |*^on" Exchange)	^""reenlnjil   I Thickening I^Flltratt
v-;   ~r~Lxr:  rTE^-irT^,  JT
                                          Sodium
                                          r/-iin1f*t i fin
  U
Waste
I Treal-mcnt
                                                    ~1
                                         I Barium
                                          Load inff

I
! Se 1 1 1
1



ing |
|

- U
1
1 Rinsing
1 	 l

_




lfuca^_Water
I
1
ent
Ing
f
Hng |
f
rlfu- LJ
>



Miing
_J
k
 f Drying
J

Pelletizlnj
* f-

                                                                        Col
                                                                                    Water
                                                   Makeup
                                                   Barite
                                                     -*[ Cooling')

                                                        ~
                                                                                Milling
                                                                               Leaching
                                                               Nat.
                                                               Gas
                                                                                   Ifu-
                                                I    ,	CZ
                                                I  UI Dilution
                                                       >    1
                                                                 .Water
        Flue Gas
          B\ -Product
          I ecoverv
                                                                        Barium Conversion
 Waste
Treatment
       Resin
     Regeneration
                                                     Effluent
                                                    Treatment
  Federal \af er qia 1 ity
  Admin 1st rat ion
  Contract No.  14-12-'>1
            FIGURE - 22

       : OCK DIAGRAM FOR AMD

        NKITRA'TSI'LFATINT
                                                       Catalytic, Inc.
                                                       Cont . 41520
                                                       Dec., 1969

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to form in the system.  Even with such an aid, slow, hindered
settling is expected to prevail; however, the special design
of the pretreatment clarifier is expected to allow relatively
complete settling of the iron-aluminum floes; removal of this
sludge will be aided by rake type agitator arms which act as
a centrifuge to further concentrate the slurry in a bottom
center well.  Due to the rather non-compressible settling
characteristics of the sludge, a very dilute underflow is
expected ranging from 1-15 percent solids, depending on the
effectiveness of the anti-flocculent and the rake arms.  There-
fore, dewatering of the sludge is very important in order to
recover the valuable AMD, and also to hold to a minimum the cost
of sludge disposal.  Thus, the underflow slurry is pumped to a
high speed disc-bowl centrifuge which will recover most of the
AMD water, sending it back to the clarifier feed zone.  The
partially dewatered sludge containing up to 20-50 weight percent
solids is then mechanically conveyed to disposal barrels or tank
cars, from whence it will be either sealed in mines or deep-
welled below the river bed.  The centrifuge contains a manually
adjustable batching nozzle and pulse control which will maximize
the sludge dewatering operation.

Removal of Sulfates and Carbonates from AMD

Pretreated, neutralized mine drainage from the overflow weir of
the pretreatment clarifier is pumped to the pressurized desulfating
section of the production plant.  The feed is distributed in
parallel to four barium exchangers.  Barium loaded ion exchange
resin is hydraulically conveyed from the regeneration area to three
resin screens in parallel, which in turn send the resin retains to
each of the first three exchangers.  The resin feed lines to the
reactors are equipped with rotary vane feeders with preset timer
control.  Plant air is available to equalize reactor pressures in
the resin feeders.  Compressed flue gas from the Barium Conversion
Area is supplied to the first exchanger at 75 psig.  Each exchanger
is equipped with relief and control systems, so that the pressure
will be gradually let down until the last exchanger is operating at
only 3 psig.

Thus, the C02 in solution will vary from 1750 ppm in the first,
down to 440 ppm in last exchanger.  The sulfate to carbonate ratio
in the reactor chain is expected to approach unity to an extent
which will allow just enough elution of barium from the resin to
incur complete desulfating of the AMD.  Thus, the parallel feed
                               -70-

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arrangement and pressurized system are both seen as excellent means
of monitoring the barium ion concentration in the AMD; this is
important in both preventing the precipitation of barium sulfate
in the resin beads and fostering its precipitation from the AMD.

Cascading from one barium exchanger to the next, the total flow
leaves the last vessel, flowing by gravity to the primary screens
where the resin is retained and the AMD, containing barium sulfate
fines, is pumped to the degasifier column.  The AMD enters the top
or fifth tray of the "A" Section of the column; air is blown from
the "A" Section bottom, contacting and freeing the dissolved C02
The AMD-sulfate containing barium then passes to the bottom or "B"
Section of the column; concentrated sulfuric acid is metered to the
top tray (20th) of this section and reacts with the carbonates which
formed during the ion exchange cycle.  C02 is evolved and is con-
tacted with air from a blower.  Thus, the degasified and decarbonated
AMD leaves the bottom of the column and is pumped to the terminating
reactor.  The total CC>2 and air leaves the top of the "A" Section to
atmosphere.

In the terminating reactor, AMD containing both precipitated and
some dissolved barium is mixed with pretreated AMD containing
sulfates which react with barium to form a precipitate.  Both barium
sulfate formation and crystal growth are fostered in the irreversible,
and therefore terminating steps of the reaction model (see page 11).

The AMD - sulfate slurry is then gravity fed from the terminating
reactor overflow weir to the thickener.  Flocculent is introduced
along with the slurry to the feed reaction zone of the clarifier,
which promotes agglomeration of the fines.  Clarified AMD, which
is essentially free of sulfates and carbonates, overflows the weir
and is collected in a surge tank for pumping to the product filter.

Underflow, consisting of 20 to 40 weight percent barium sulfate
suspended in AMD, is pumped to a centrifuge where it is dewatered,
resulting in a cake containing 20 to 50 percent water, which is
conveyed to the barium conversion area.

The sulfate clarifier overflow may contain up to 100 - 200 ppm
turbidity as barium sulfate fines.  This shall be removed in a
product filter system containing precoated cartridges in a
continuous cycle.  The filter is equipped for alternating back-
washing-precoating, and on-line operations.
                               -71-

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Pretreated, desulfated, decarbonated and filtered product AMD
emerges from the filter.

Resin Regeneration

The ion exchange resin that was separated by screening from the
desulfated AMD containing barium sulfate fines, in the primary
filter, is partially exhausted, consisting of sodium, calcium,
magnesium, and mostly of barium ions.  The resin falls from the
screen into an agitated ion exchanger where a 0.1N sodium chloride
solution is introduced.  Due to an excess of sodium chloride present
in the system and with equilibrium conditions prevailing, some
barium will be eluted from the partially exhausted resin.  However,
because of the selectivity of the resin toward barium, it is
expected that most of the calcium and magnesium will be eluted
rather than the barium.  Also, a small amount of the desulfated
AMD, containing unreacted barium, from the degasifier feed pump
may be introduced to prevent barium elution by the salt.  The
resin-salt solution slurry overflows the exchanger weir into a
secondary screen which passes the waste stream, while retaining
the resin; the latter is transferred by bucket elevator to the
resin loading section.  Resin is passed to one of two moving bed
loading columns.  Each column consists of internal fiberglass tube
bundles and distributors to prevent bridging and plugging of the
resin beds.  The resin falls by gravity counter-current to ascending
barium containing liquor.

The barium loaded resin is collected in the bottom sections of the
loading columns.  The streams are then combined and flow into the
rinsing column which also consists of an internal fiberglass tube
bundle with distribution and collection system.  The resin falls
by gravity in counter-current flow to the ascending process water
which removes any entrained regenerant from the resin pores.  Rinsed,
barium loaded resin leaves the bottom of the rinse column and is
passed through the suction of an eductor and is hydraulically conveyed
by 100 psi process water to the resin feeding system screens in the
desulfating section.

The regenerant solution is produced in the barium conversion area
and is sent first to a feed tank where it is diluted with process
and rinse water to the proper barium concentration (O.lN) from
whence it is pumped to the parallel loading columns.
                               -72-

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Effluent Treating

The effluent from the top  (clarification) sections of the loading
columns consists mainly of dissolved sulfides and hydrosulfides
of sodium with traces of barium, calcium, and magnesium ions.
The solution flows by gravity, into an agitated treating tank where
some pretreated AMD is also introduced.  The sulfate in the AMD
reacts with the traces of dissolved barium to form an insoluble
precipitate.  The slurry is pumped to a separator in the carbonation
area where the precipitate is removed by gravity settling; the clear
effluent overflows the weir and is pumped to the carbonation column.
The barium sulfate pulp is drained from the separator bottom and is
pumped to a centrifuge which dewaters the slurry generally from
50 to 80 percent solids.  The cake discharge is then conveyed to the
barium conversion section, the liquid filtrate being returned to
the separator or transfer pump suction line.

Waste Treatment

The waste stream from the secondary resin screen in the resin
purification area of the regeneration section contains mainly a
dilute solution of sodium, calcium, and magnesium chlorides, with
traces of barium.  In order to remove barium in the form of sulfate,
the waste stream flows by gravity into an agitated tank where it is
mixed with pretreated AMD containing dissolved sulfate.  The waste
stream containing barium sulfate fines is passed into a waste
separator which removes a clear, barium free waste stream as over-
flow to drain, and a barium sulfate slurry as underflow, which is
returned to the sulfate centrifuge for dewatering.

Barium Sulfate Conversion

Pulverized coal is received in hopper bottom cars and is unloaded
and conveyed to conical bottom storage bins.  Pulverized barites
are also received and stored for barium makeup.  The coal is
conveyed from storage to a feed hopper, while barium sulfate slurry
is simultaneously conveyed from the sulfate centrifuge and the
sulfate recycle centrifuge (unreacted sulfate from the leaching
operation) to another feed hopper where barite makeup is also added.
Both bins are equipped with vibrators, feeders, and controls neces-
sary to maintain a preset, ratio of coal to sulfate into a geared
mixer which blends the two constituents.  The mixture is then hoisted
to a mixed sulfate surge bin which feeds the roasting system.
                               -73-

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Initially, the net barium sulfate-carbon mixture is dried to
5-10 percent moisture in a direct heat rotary dryer which
uses hot flue gas from the calciner in counter-current flow
as a drying medium (an auxiliary burner is also provided).
The relatively dry mixture is then fed to a rotary drum
pelletizer.  The barium sulfate, which has a particle size
averaging one micron, and the pulverized coal, passing
through a 100 mesh screen, are balled to pellets or modules
by the rotating action of the pelletizer surface.

The pellets enter the calciner which is a direct heat, brick-
lined rotary kiln, fired with natural gas in counter-current
flow; controls at the firing end maintain a reducing atmosphere.
The intimately mixed coal reacts with the barium sulfate at a
temperature of about 2000F to produce barium sulfide at a
conversion rate of 92 percent.  Because the barium sulfide is
highly reactive at elevated temperatures and is rapidly oxidized
to the sulfate form in the presence of oxygen (air), the calcine
must be cooled below 200F before grinding.  Cooling is effected
in an indirect water spray rotary cooler which utilizes process
water at 77F.

The hot flue gases from the kiln are passed through the rotary
dryer, where they are cooled below 200F and contain approximately
20 percent (vol.) carbon dioxide.  Part of this is blown to the
carbonation section of the plant to be described later.  The other
portion of the flue gas is compressed to 100 psig and is sent to
the pressurized exchangers in the desulfating section.  Cyclones
are included with both the calciner and dryer to eliminate entrain-
ment in the flue gas.

The cooled pellets are then conveyed to a calcine mill which grinds
the pellets, which consist of approximately 88 percent barium
sulfide, 10.5 percent barium sulfate and 1.5 percent ash, to about
20 microns fineness at 176F.  The mill includes equipment for
classification and dust collection.

The calcine fines are conveyed to a surge bin which feeds into the
first reactor of two in series, which are part of a two-stage
counter-current leaching operation.  Process water is preheated in
the calcine cooler and combined with heated wash water from the
sulfate recycle centrifuge.  This combined stream is then passed
                               -74-

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into the Second Stage Contactor which dissolves any entrained
sulfides which are in the First Stage Thickener underflow.
The slurry then overflows into the Second Stage Thickener
which separates the undissolved sulfate and ash by gravity
settling; the underflow is pumped to the sulfate recycle
centrifuge which dewaters the sulfate.  Approximately a 50 - 70
percent  (by wt.) discharge cake is conveyed, along with the
barium sulfate from the dewatering section of the desulfating
process, back to the feed hopper for mixing with coal.  The filtrate
is combined with the feed to the Second Stage Contactor.  The
Second Stage Thickener overflow, consisting of a dilute barium
liquor,  is then heated from 118 to 156F and is passed into the
First Stage Primary Contactor where it is mixed with the calcined
fines from the mill section; the latter is fed at a preset ratio
to the amount of process water which is fed to the Second Stage
Contactor.  The partially dissolved slurry overflows into the
First Stage Secondary Contactor, where most of the remaining
barium sulfide is dissolved.  The slurry, consisting of barium
sulfate, ash, and some undissolved sulfide entrainment, with the
sulfide  solution is gravity fed to the First Stage Thickener.
The sulfate, ash and entrained sulfide settles out and the underflow
is pumped to the Second Stage Contactor where the remaining traces
of sulfide is dissolved in the first stage feedwater.  The full
miscella (barium containing liquor) from the First Stage Thickener
overflow is pumped to the Feed Dilution Tank in the resin
regeneration area.

Following centrifugation and washing, part of the sludge from the
recycle  centrifuge is discarded (blowdown line) to keep down the
ash concentration.

Carbonation

The treated effluent from the loading columns in the Resin Regenera-
tion section is pumped to a separator which pumps the barium sulfate
underflow to a centrifuge for dewatering.  The separator overflow,
consisting of less than 0.1N sodium sulfide with trace barium, is
pumped to the Carbonation Column in which it flows by gravity from
the top  stage (tenth) to the bottom section in which flue gas is
fed counter-currently.  Each stage consists of a sieve tray arrange-
ment with a paddle-type agitator.  The flue gas is blown from the
roasting section of the barium conversion loop and contains 20 percent
(by vol.) CC>2.  The carbon dioxide contacts the sodium salt solution
                               -75-

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and reacts to form sodium bicarbonate.  Hydrogen sulfide is
formed during the reaction and is passed out the top along
with the flue gas.  This is blown to the sulfur recovery
plant which converts the H^S to elemental sulfur which is
stored and sold as a by-product.  The sodium bicarbonate solution
is returned to the pretreatment section where it is used in the
neutralization of AMD.  Any excess of sodium bicarbonate will be
discarded.

Sulfur Recovery

The sulfur recovery plant will be purchased as a package unit
and will recover 1.013 tons/day of sulfur (elemental).  Flue gas
containing from 10-20 w/o hydrogen sulfide will be supplied to the
plant at a rate of 100-150 CFM.

Recovery of Salable By-Products and Logistics

By-Product Recovery

Four by-products could, in theory, be produced in the operation of
the AMD Neutradesulfating Production Plant:

1.  Sodium bicarbonate solution (0.7 percent by wt.) can be produced
    in excess over that needed for pretreatment.  This would yield
    about 2 - 2.5 tons/day solids.  However, the cost of recovery
    and purification of this product is prohibitive for yields in
    this range.

2.  Sludge from the pretreatment area, consisting of iron oxides,
    aluminum hydroxides, hydrates of both, and calcium carbonates,
    could be dried to yield 1.6 ton/day.  However, to date there
    is no market value for such a mixture, and separation techniques
    could not be justified from both an operational and economic
    basis.

3.  Sulfur will be recovered by a packaged process which converts
    H2S in flue gas into elemental sulfur at a rate of 1.01 ton/day,
    which is essentially pure and has market value.

4.  A waste stream from the sodium completion step of the resin
    regeneration section, which has been treated to remove any
    dissolved barium, overflows from the waste treatment clarifier
    into a drain.  It consists of dilute (1%) quantities of sodium,
    calcium and magnesium chlorides.  Cost of separating and recover-
    ing of individual components are prohibitive and impractical.
    The stream, as is, has no known market value or use.
                               -76-

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Logistics Associated with the Neutradesulfating Process

In a 1MM GPD neutradesulfating production plant, the relatively
low rates of by-product accumulation and chemical consumption require
only one railroad car per two weeks which would carry all the
chemicals necessary to maintain a full production rate for that
period.  Sulfur being the only marketable by-product will need but
one car every three months.

Therefore, trucking the materials seems to be a practical alternative.
A typical self-unloading pneumatic bulk trailer will carry 15 - 20 tons
of sulfur.  Ordinary dump trucks with tight-fitting tarpaulin covers
could also be used if the sulfur is bagged.  Only one truck is needed
per two-week period for the by-product; whereas three trucks per two-
week period would be required for the entering chemicals.  Catalytic's
design has provided for two-week storage capacity.

An initial supply of process water is required for startup, which shall
be produced thereafter by the process.  Therefore, a 20,000 gallon
storage pond for process water requirements has been included (10* x
20' x 2').  Also, an initial 1,000 cubic feet of resin will be required.

Table XVII describes the trucking (or R.R.) traffic for the plant.
Table XVIII lists utilities for the production plant.
                               -77-

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                                TABLE XVII
Production Plant Daily Traffic Schedule
1. Products (leaving)
Sulfur
Sludge (*}
(**)
Waste Stream
2. Chemicals (entering)
Coal
Salt Cake
Barites
Sulfuric Acid
(42 gal. bbl.)
Flocculent
Resin
Production Rate Transportation
of Usage By R.R. or By Truck
Ton #
Tons /Day Tons /Car # Cars Truck Truck

1.013 50 - 70 1/3 Mo. 15 - 20 1/2 Wk.
1.608 - deep-welled -
3.26 - discarded

0.9 50 - 100 - 20-30 1/2 Wk.
3.636 70 - 100 1/2 Wk. 30 - 40 2/2 Wk.
0.12 150 - 180 - 60-80 1/Yr.
0.3 60 - 120 - 25-35 1/3 Mo.
0.084 35-70 - 15 - 20 1/6 Mo.
0.534 60-70
(per yr.)
4.742 Ton 70 - 100 1 car/2 Wk. 3 Truck
Day 2 Wk.
 (*)  Dry Basis - Sludge to be welled as 507,, water slurry.
(**)  Dry Basis - Exits in a 0.5% solution by weight.

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                 TABLE XVIII
               Production Plant
               Utilities Summary
                       Flow          Condition



Steam               100 Lbs/hr.      @ 5 psi

                                         Sat'd.
                           (*)
Process Water       240 GPM           25 psig
Cooling Water        50 GPM           25 psig






Electricity         850 KW





Plant Air           100 SCFM        @ 80 psi
(*) Total process water input.
                     -79-

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                          PLANT EQUIPMENT
The production plant is a conceptualized design for a 1MM GPD commer-
cial unit.  All sizes are shown on the Appendix D drawings R-211 and
R.-212, which include production arid waste recovery sections, respec-
tively.  Cost estimate for a 1MM GPD unit was based on the prices of
equipment shown on these drawings and described as follows:

Furnace

A natural gas fired packaged boiler producing saturated steam au 20 psig
with a capacity of 100,000 BTU/hr. shall be furnished.

Reactors
All reactors shall be tar modified epoxy-coated or rubber-lined carbon
steel vessels, having ASME dished heads, and top entering, center
mounted, variable speed, turbine-type, rubber-coated steel agitators.

Towers

External tar modified epoxy-coated carbon steel with fiberglass dis-
tribution system and internals are specified for the moving bed Resin
Regeneration and Rinsing Columns.

The Degasifier Column is divided into two sections; the top section has
a 5-stage sieve tray, 1-pass arrangement and the bottom section has 20
stages with a 2-pass, sieve, tray design.  Each has an air distributor
s y s t em.

The Carbonation Column is 10 stages with a turbine paddle agitator on
each perforated tray^ running from a central shaft.  These towers are
of epoxy-coated carbon steel.

Heat Exchangers

One rotating solids cooler and two conventional shell and tube heaters
are of carbon steel construction.
                               -80-

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Vessels

In general, 316 stainless steel or polypropylene are specified for
chemical service, rubber-lined or epoxy-coated carbon steel for acid
service, epoxy-reinforced fiberglass, or teflon lining, for resin
conveying and feed tanks, and carbon steel for storage.  Exceptions
include monel, and high silicon iron for special chemical and waste
applications.
Centrifugal pumps are specified for AMD and most chemical services,
and are generally of stainless steel with wetted parts.  Metering
pumps for special chemical feed systems have stainless steel wetted
parts.  Diaphragm pumps are specified for many of the solids handling
applications and are hypalon lined, with hypalon diaphragm and quick
opening valves.

Materials Handling

Sludge, sulfate, and coal conveyors are either the belt, flight, or
tubular type.  Rubber belts with moulded rubber cleats, dust-free
totally enclosed drag flights, and carbon steel with stainless steel
links and neoprene flight material shall be used.  Coal storage bins
and conveyor feed lines shall include flanged discharge openings,
terminals, anti-friction head bearings, take-ups, and drip-type
lubricating systems.

Sulfate and carbon feeders shall be of carbon steel and are complete
with vibrating mechanism and supports.

A 40-foot bucket elevator with lexan buckets, plexiglass inspection
windows, and mild steel housing with acid resistant coating shall be
used for resin service; and an inclined 43-foot vertical, 5-foot
horizontal travel skip hoist elevator of epoxy-coated carbon steel shall
be used for calcination feed.

A 20-foot vertical, 5-foot horizontal inclined bucket elevator of epoxy-
coated carbon  steel shall be used for the leaching feed system.

A hydraulic conveying system, which includes teflon coated  lines, and
PVC eductor are used to transfer resin feed.
                               -81-

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The Sulfate-Carbon Mixer 'consists of dual, stainless steel agitator
paddles with a retention and dam at discharge end and epoxy-coated
steel trough.

The Sulfate-Carbon Dryer is a counterflow rotary self-supported drum
with auxiliary steam heating coil, and is of epoxy-coated steel con-
struction.

The Pelletizer shall be a rotary-coated carbon steel unit with stain-
less steel interior surface.

The Calciner shall be a natural gas-fired rotary kiln of brick-lined
carbon steel, equipped with inlet and outlet feed, natural gas, air,
and flue rotary connections.

The Calcine Mill is a comminuting machine with 316 stainless steel
chamber and 414 or 420 hardened, tempered stainless steel hammer
blades, with stainless screens.  This unit also includes a stainless
steel collector and cyclone, cast iron vacuum producer-  and a negative
pressure rotary air lock.

All of the above mentioned equipment are for continuous  service, and
includes epoxy-coated or acid-resistant painted surfaces.

Filters, Screens, Centrifuges, and Other Separators

Clarifiers shall be field erected and are complete with  central inlet
feed chambers, rakes, motorized driver, controls, alarms, walkways,
etc.  All wetted parts are to be tar-modified epoxy-coated carbon steel.
The clarifiers generally include counterflow detention hoods, orifice
holes, and collection flumes for optimum clarity effluents; also, Neptune
60 inclined tube modules are provided to increase collection capacity.

The Product Water Filter shall consist of two carbon steel units each
containing 260 disposable cartridges, of polypropylene media and matrex
on 304 stainless steel cores, with a 1 micron rating. One unit shall
be backwashed and precoated while the other is in service.  Pressure
drop measuring devices and manually operated valves are  included.

Vibrating screens are used to separate resin from various slurry streams.
All units will be furnished with plastic-coated vibration eliminator
springs, epoxy-coated externals, 316 stainless steel wetted parts, and
neoprene connectors on discharge spouts.  All are single stage except
the Exhausted Resin Secondary Screen which has an additional deck for
removal of attrited resin beads.
                               -82-

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The Traveling Water Screen consists of a belt of inclined screening
panels with clear opening stainless steel wire mesh, and bottom
ledgers to retain floating and suspended matter.  Jet sprays backwash
the material into a trough or disposal sluice.

The centrifuges are continuous horizontal model decanters of various
horsepowers and bowl sizes, equipped with motor, starter, and vibration
isolators.  The exception is the sludge centrifuge which contains a
continuous vertically wheeling basket from which the sludge will be
periodically skimmed by timer control and is equipped with spring
mount and flexible connections.  All units shall be of 316 stainless
steel wetted parts and epoxy or chemically coated externals.

Agitators

Generally, top entering, center-mounted units of rubber-covered or
epoxy-coated carbon steel will be used.  Reactor units will have dual
turbine blades with stuffing boxes, 8-inch flanges for top center
mounting, and stabilizing rings.

Smaller units with less vigorous mixing requirements and open tank
operations shall be portable propeller type.  Chemical dissolution
service will use stainless steel units with high powered center-
mounted propeller blades.

Packaged Units

The Instrument Air Compressor consists of a centrifugal drive to
deliver oil free air at 100 psig, a vertical air receiver, and air
dryer.  The Flue Gas Compressor is of similar design but includes
an aftercooler and a special bronze pump and trim.

Centrifugal blowers will be used for supplying air to degasifier and
kiln, and shall be of tar-modified epoxy-coated carbon steel.  A
stainless steel blower will be used to transfer H^S containing flue
gas to the Recovery Section.

A Salt Lixator, consisting of a concrete cubicle divided into two
sections, supplies a concentrated brine solution to the process.
Rock salt is dumped in the loading hatch of either cubicle in which
a water level is maintained.  The solution is collected in a brine
well for pumping.  Each cubicle can be drained and cleaned independently.
                               -83-

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                             LAYOUT

A conceptualized plan for a 1MM GPD commercial unit is sketched in
dwg. B-211 Appendix D.  This includes estimated sizes for offsite
considerations, and includes two-week storage allowances for incoming
chemicals.  Included are warehouse, maintenance, operating, and
administration spacing considerations.  Also, a 20,000 gal. pond
is included for storage of process water for startup requirements.
Equipment is grouped in block form and attention is given to both accessa-
bility and proximity, which at the same time, attempts to hold materials
handling costs to a minimum.

The sketch assumes that this location relative to the river side is
possible; it is also assumed that railroad and roadway routings could be
oriented in the illustrated fashion.  However, there is a built-in
flexibility in the blocked plan that would allow for easy manipulation
of all vital units.  The sulfur recovery plant space is an estimate due
to lack of available sizing data for the packaged unit in this relatively
low capacity range.

Approximate overall dimensions of the land required for layout of the 1MM
GPD unit are 350' x 200', or 1.6 acres.
                              - 84 -

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                              ECONOMICS
Capital Investment

The estimate for the capital investment of the neutradesulfating plant
and by-products facilities is based on vendors'  written or verbal quota-
tions for the major equipment.  The neutradesulfating plant has been
regarded as an independent "grass roots" plant including storage ware-
houses, offices, laboratory, etc.  A 14-day storage capacity for chemicals
and by-products is included.

The total capital investment for the 1MM GPD plant of $4,956,600 is based
on a major equipment cost of $1,563,200 and was  built up to include all
construction cost including a 15% contingency and a working capital of
10%.  No charges were made for land.  Based on this, the total capital
investment for 10, 25, and 100MM GPD were calculated at a prorating
factor of six-tenths.  A summary of investment capital is presented in
Table XIX.

Operating Costs

The operating costs were calculated for an operating year of 328 days
and are presented below on both a unit cost basis and an overall annual
basis as shown in Table XX, for a 1MM GPD plant.  The basis for the
various charges follows.

Direct Costs

Chemicals;

Barite makeup is taken at $30/ton.  Its consumption is dependent on the
efficiency of the sulfate clarification.  The loss due to entrainment in
the overflow is estimated as 1.57=, of the barium sulfate formed in the
desulfating ion exchange cycle or approximately 10 Ibs/hr.

Coal is used in the reduction of barite in the rotary kiln (calcination
step) at a rate of 75 Ibs/hr.  The coal is taken at a cost of $10/ton for
powdered anthracite at the plant site.

Salt (NaCl) is used to transfer the spent resin into solely the sodium
form.  A rate of 300 Ibs/hr. is needed.  The price is taken at $12/ton.
                                -85-

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                          TABLE XIX




      CAPITAL INVESTMENT FOR NEUTRADESULFATING PLANTS





1MM GPD PIANT
Purchased Equipment               $1,563,200




Fixed Capital Cost                                 $4,506,000




Working Capital                                       450.600




      Total Capital Investment                     $4,956,600




10MM GPD PLANT




Purchased Equipment               $6,200,000




Fixed Capital Cost                                $17,900,000




Working Capital                                     1.790,000




      Total Capital Investment                    $19,690,000




25MM GPD PLANT




Purchased Equipment              $10,770,000




Fixed Capital Cost                                $31,100,000




Working Capital                                     3.110.000




      Total Capital Investment                    $34,210,000




100MM GPD PLANT




Purchased Equipment              $24,700,000




Fixed Capital Cost                                $71,200,000




Working Capital                                     7,120.000




      Total Capital Investment                    $78,320,000
                              - 86 -

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                          TABLE XX

             ECONOMIC EVALUATION FOR A 1MM GPP
                  NEUTRADESULFATING PLANT
TREATING COSTS
  DIRECT COSTS
    Chemicals
$/1000 Gallon
 $/Year
Barites @ $30/ton
Coal @ $10/ton
Salt @ $12/ton
Flocculent @ $.73/lb.
Resin @ $17/CF.
Sulfuric Acid @ $27/ton
.002
.009
.044
.012
.0002
.008
        Subtotal Chemical

    Utilities
     .0752
      Steam Self-Supporting
      Process Water @ $.22/1000GAL. .074
      Electricity   @ ,0075/KWH     .147
      Fuel Gas      @ $.35/MM Btu   .008
        Subtotal Utilities

    Labor & Supervision

      Direct Labor
      Supervision
      Maintenance

        Subtotal Labor

      Repair and Maintenance
     .229
     .220
     .044
     .040

     .304

     .386
      Supplies (15% of Maintenance) .058     	

        Total Direct Cost                    1.052

INDIRECT COSTS

    Plant Overhead (50% prod, labor) .152
    Payroll Overhead (20% prod.     .061
      labor)
    Packing and Shipping (*)        .097     	

        Total Indirect Cost                  0.310
   656.
  2952.
 14432.
  3936.
    66.
  2624.

 24666.
 24272.
 48216.
  2624.

 75112.
 72160.
 14432.
 13120.

 99712.

126608.
 19024.
                         49856.
                         20008.

                         31816.
                                 345122,
                                 101680.
(*) Includes waste transport allowance for sludge
                                    - 87 -

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                          TABLE XX
TREATING COSTS
  FIXED COSTS
$71000 Gallon
    Depreciation (4.6%, 30 yrs.)    .943
    Property Tax (2% Fixed Cap./yr.) .275
    Insurance (1% Fixed Cap./hr.)   .140

        Total Fixed Cost
$/Year
                          309304.
                           90200.
                           45920.
                1.358
          445424.
        Total Treating Cost
CREDITS
    By-Product Sulfur @$30/ton

        Total Credits

        Net Unit Cost of
        Neutradesulfated Water
                2.720
          892226.
      .030
                 .030
                                              2.690
            9840.
                                    882386,
                               - 88 -

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Flocculent is used to improve settling rates in the pretreatment
and sulfate clarifiers.  A rate of 0.7 Ibs/hr. is needed (*),  but
due to the high price of $.73/lb., the relative cost is significant.

Ion Exchange Resin makeup rates of .012 CF/day are accounted at a
nominal cost of $17/cubic foot.

Sulfuric Acid is used in the bottom section of the degasifier column
to remove remaining carbonates and bicarbonates formed in the pressur-
ized desulfating unit, which cannot be removed by merely air sparging.
This amounts to 1.6 gph, at a unit cost of $27/ton of 66 Be Acid.

Utilities;

Electricity is the main contribution to the utility costs,  being
consumed at a rate of 825 KWH.  The recommended price of $.0075/KWH
is used.

Fuel Gas is a minor cost, and is taken at $.35/MM BTU.  1MM BTU/hr.
is required for the preheat and calcination of barium sulfate.

Process water consumption is due to loss of water in outgoing waste
streams amounting to 83 GPM; water cost is taken as $.22/1000 gal.

Labor and Supervision;

A total of eight operators ($9,000/yr.) is needed for a 7 day, 40
hour, 3 shift/day, production week.  One supervisor ($14,000/yr.),
two maintenance men ($6,500/yr.) and two lab technicians ($8,000/yr.)
are on a 40 hour week; administration includes a plant director
($15,000/yr.), technical advisor ($12,000/yr.), and a secretary ($6,500/yr.)
all of whom are salaried.  Total labor and supervision comes to 11.4% of
manufacturing cost.
(*) Based on 1 Ib/ton dry solids
                              - 89 -

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Repair and Maintenance Charges were estimated from average main-
tenance costs of similar equipment from past applications.  These
were updated using a maintenance index.  This amounts to 2.81% of
fixed capital cost.  Labor due to special maintenance problems is
a part of this charge (usually about 50% of repairs including 10%
supervisory labor).

Supplies are estimated as 15% of maintenance and repairs.

Indirect Costs:

Plant overhead is  taken as 50% of the productive labor cost, with a
20% of productive  labor allowance for payroll overhead.  Productive
labor includes direct labor, supervision, and maintenance.

Packing and shipping charges are taken as $1.00/ton for sulfur, and
.03c/lb. trucking  allowance is made for waste (sludge) disposal.

Fixed Charges:

Depreciation charges are amortized at 4.6% for 30-year expected
life.

Property Tax and Insurance are taken at 2% and 1% of fixed capital
investment per year respectively.

By-Product Sale

Sulfur is the only by-product that will be upgraded for sale.  All
the barium sulfate produced is used in the conversion and regene-
ration steps.

a.  The pretreatment sludge and the waste stream have no realizable
    market value at the present time.

b.  Though the neutradesulfating process does generate other materials
    with a soluble potential, sulfur is the only one that has to be
    treated in one way or another to reduce its strong polluting
    effect.  At the same time, the reduction of hydrogen sulfide to
    sulfur is a well established process that will require an invest-
    ment lower than that required by any treating system.

The 1MM GPD plant  will produce close to one ton of sulfur per day that
is rated at $30/ton.
                               -90-

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Cost of Neutradesulfating AMD

The net unit treating cost sums up to $2.69 per thousand gals, of
feed.  This is equally correct with regard to 1000.gal. of treated
water since an amount of water almost equal to the total purchased
process water (added to product stream) leaves in subsequent waste
streams.

Expansion of Cost of Neutradesulfating for Larger Capacity Plants

Contrary to the economic factors which were designated in the
ground rules for the proposal (7), as stated in the introduction,
the true expected life of equipment is more likely to fall in the
10-20 year range.  Therefore, in Figure 23 the unit cost of neutra-
desulfating is shown as a function of plant capacity for 10, 20, and
30 year expected life spans.  These costs are relatively high for
capacities in the 1MM GPB. range, but no significant change is
noticed beyond the 25 MM GPD capacity.  Figure 24 shows the effect
of investment on the unit cost for 15 and 30 year lives.  These
figures are based on the simplified equation:

           MC   =   0.45 + (.0773 + 1.37 erf (4.6-n)x UFC
                                                      SO.4

           MC   =   Manufacturing Cost in $/1000 Gal.

           erf  =   Capital Recovery Factor for 4.6% Interest
                     over n years of life (Amortization)

           UFC  =   Unit Fixed Capital investment for 1MM GPD
                     plant in $/1000 gal/yr.

           S    =   Plant capacity in MM GPD.

           n    =   Expected life, years.

The assumptions are those stated in the Operating Cost Section;  total
labor and supervision was taken as 11.47o of manufacturing cost,  and
repair and maintenance is assumed to be 2.870 of fixed capital per year.
These figures were calculated from the 1MM GPD treating costs from
Table XX.  A 10% allowance is made for working capital
                                 - 91  -

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O

O
o
O
en
O
                        Note: n = Expected  life of  plant  in yeai
   1.0 -H
                              40          60


                          Plant Capacity - MMGPD
100
         Figure 23:   Effect of Plant Capacity on the Un.t Cos;

                     of NeutradcsulfaLing
                                          CCCo. Job -V1520

                                          U. S. Dr-pt . of

                                               Conr. 1'4-12-518
                              -92-

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The curves of Figure 23 show that the cost for a 1MM GPD plant is
relatively high, falling sharply as the plant capacity increases
to 10MM GPD, then sloping off until no significant change occurs
beyond the 25MM GPD capacity.

From Figure 24, it is noticed that capital investment has a great
effect on the unit cost of a 1MM GPD plant, but has a much decreased
effect on the 10 and 25MM GPD units.  If capital investment were
only 50% of that realized, the cost would be about $0.80/1000 gal.
for a 10MM GPD plant with either 15 or 30 year lives.

However, if a 1MM GPD plant could be built for only half the
estimated capital investment, the manufacturing cost would still
be a relatively high $1.82/1000 gal. for a 15 year life, and
$1.57/1000 gal. if a 30 year life were realized.

It can be seen from both figures that for the estimated base capital
($4,506,000), the 25MM GPD plant would be a good cut point with
only a small saving realized after this point for additional capacity.

Economic Feasibility of Neutradesulfating Plant
                                             (it)
Unless fixed capital could be sharply reduced; ' the 1MM GPD plant
would operate at a cost which is three times higher than some of the
processes which have been reported (8).  However, consideration
should be given to a plant in the 10-25MM GPD range for which the
economics are more favorable.

Future research developments and waste product marketability could
shed new light on the subject.  However, at the present time, the
neutradesulfating of acid mine drainage waters would dictate the use
of 10-25MM GPD plants at unit costs which would not likely be below
$1.00/1000 gal. for a 30 year write-off.
(*) See page ix for new approaches to basic process, which offer
    potentially lower treatment costs.
                               -93-

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D/AGRAM OF A I MM GPD
                AC/O

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J>/AGAM 0F A I MM &PP PRODUCTION
            96

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           D/AGZAM Of A I MM GPP  PRODUCTION

AJEUJR.APESULFATIN6 0f AC/0 MINE
                       
3

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                                                     Too GSM * u> rt* as f
                  J>/AGRAM OF A I MM GPP  PROPUCTfW  PLANT
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\
                                                                                                                                        FEDERAL WATER QUALITY ADMINISTRATION
                                                                                                                                            DEPARTMENT OF THE INTERIOR
                                                                                                                                         CATALYTIC  CONSTRUCTION  COMPANY
                                                                                                                                                                   LAYOUT
                                                                                                                                          I MM 6PD PRODUCTION PLANT
                                                                                                                                                         O OOP Mint ML*
                                                                                                                                        V 41520
B-2

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1

5
.Access ion Number
Q Subject Field &. Group
05 D
SELECTED WATER RESOURCES ABSTRACTS
INPUT TRANSACTION FORM
Organization
        Catalytic, Inc., Philadelphia,  Pennsylvania
    Title
        Neutradesulfating Treatment Process for Acid Mine Drainage
 10
    Authors)
        Engineering Division
16
Project Designation

     14010 DYH
21  N
 22
    Citation
        Federal Water Quality Administration, Department of  Interior,  Report No.
 23
    Descriptors (Starred First)
        * Resins, * Anion  Exchange, *Ion Exchange
 25
    Identifiers (Starred First)
        *Barium Elution,  *Carbon Dioxide Pressure
 27
    Abstract
        A process has  been developed for treatment  and  removal of major pollutants
        in surface  streams of Appalachia caused by  acid mine drainage.  The raw water
        is neutralized with sodium bicarbonate to precipitate iron and aluminum,
        followed by cation exchange to remove sulfate.  Barium is eluted from the
        exchange resin and reacts with sulfate in the water to form a precipitate.
        Barium is recovered from the precipitate and is processed to rechange the
        exchange resin.  The water is further treated to  remove hydrogen sulfide
        by conversion  to sulfur as a saleable byproduct.   The process minimizes
        the cost of sludge and waste disposal while offering a product water meet-
        ing the highest interstate water quality standards.  Treatment costs are
        $2.69 per thousand gallons for a 1 million  GPD  plant.  The project was
        terminated  at  the end of Phase I due to the high  estimated cost of treat-
        ment .
Abstractor
        H.H. Bulkowski
                               Institution
     Catalytic, Inc.
 WR;t02  tREV. JULY 1969)
 WRSI C
           SEND TO: WATER RESOURCES SCIENTIFIC INFORMATION CENTER
                  U.S. DEPARTMENT OF THE INTERIOR
                  WASHINGTON. D. C. 20240
                                                            AU.S. GOVERNMENT PRINTING OFFICE: 1972 484-483/56 1-3

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