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International Journal of Greenhouse Gas Control 5 (2011) 1319–1328
Contents lists available at ScienceDirect
International Journal of Greenhouse Gas Control
journa l homepage: www.e lsev ier .com/ locate / i jggc
preliminary cost and engineering estimate for desalinating produced formationater associated with carbon dioxide capture and storage
.L. Bourciera,∗, T.J. Wolerya, T. Wolfeb, C. Haussmannc, T.A. Buschecka, R.D. Ainesa
Lawrence Livermore National Laboratory, 7000 East Avenue, Livermore, CA 94550, United StatesPerLorica Inc., P.O. Box 190, Rough and Ready, CA 95975, United StatesWater Systems Specialists, Inc., 5808 Princeton Ave NE, Seattle, WA 98105, United States
r t i c l e i n f o
rticle history:eceived 3 August 2010eceived in revised form 13 May 2011ccepted 1 June 2011vailable online 6 July 2011
eywords:
a b s t r a c t
The risk associated with storage of carbon dioxide in the subsurface can be reduced by removal of acomparable volume of existing brines (e.g. Buscheck et al., 2011). In order to avoid high costs for disposal,the brines should be processed into useful forms such as fresh and low-hardness water. We have carriedout a cost analysis of treatment of typical subsurface saline waters found in sedimentary basins, comparedwith conventional seawater desalination. We have also accounted for some cost savings by utilizationof potential well-head pressures at brine production wells, which may be present in some fields due to
arbon capture and storageesalinationrineseverse osmosissmotic pressureroduced waters
CO2 injection, to drive desalination using reverse osmosis. Predicted desalination costs for brines havingsalinities equal to seawater are about half the cost of conventional seawater desalination when we assumethe energy can be obtained from excess pressure at the well head. These costs range from 32 to 40¢ per m3
permeate produced. Without well-head energy recovery, the costs are from 60 to 80¢ per m3 permeate.These costs do not include the cost of any brine production or brine reinjection wells, or pipelines to thewell field, or other site-dependent factors.
. Introduction
A major risk associated with geologic sequestration of carbonioxide is the buildup of pressure in the subsurface due to injec-ion. The maximum sustainable pressure is limited by the abilityf the cap-rock to contain the CO2 as a low permeability barrier,nd also the potential for the overpressure to create new frac-ures, or to reactivate existing fractures, which would then openew flow paths for CO2 escape (Rutqvist et al., 2007). Moreover,mplacement of CO2 could drive the existing brine into usefulquifers where potable water would be contaminated, and couldlso increase the likelihood of induced seismicity (Bachu, 2008).
A method to avoid these potential problems is to withdrawhe saline fluid existing in the subsurface and thus reduce themount of overpressure. However, the outstanding problem withhis approach is the disposition of the withdrawn fluids. Because ofheir high salinities (>10,000 mg/L TDS), these fluids are generallyot useful for either domestic or agricultural use. The large volume
f CO2 to be emplaced implies a similarly large fluid volume muste produced. Such a large volume of brine cannot be disposed ofs-is without significant cost.∗ Corresponding author at: L-184, Lawrence Livermore National Laboratory, Liv-rmore, CA 94550, United States. Tel.: +1 925 422 9885; mobile: +1 925 667 7165.
E-mail addresses: [email protected], [email protected] (W.L. Bourcier).
750-5836/$ – see front matter © 2011 Elsevier Ltd. All rights reserved.oi:10.1016/j.ijggc.2011.06.001
© 2011 Elsevier Ltd. All rights reserved.
Desalinating the produced brines using membrane-basedtechnologies are a potential solution that would allow brine with-drawal, with disposal partly accounted for by production of usefullow salinity water. Moreover, some of the energy needed to drivedesalination could be obtained directly from the overpressurepresent in the subsurface generated by the emplacement of CO2 forsome sequestration sites. This would allow a re-capture of somefraction of the energy used to pressurize and inject CO2 into thesubsurface to be used to desalinate the produced fluids. This can bedone directly using reverse osmosis (RO), a technology that uses apressure gradient across a semi-permeable membrane to producedesalinated water.
An additional benefit is that strategic withdrawal and reinjec-tion of brines may allow manipulation of the location and migrationof the CO2 plume in the subsurface. This could be accomplishedby using brine injection/withdrawal to alter the subsurface pres-sure field so as to affect the pressure gradient surrounding theCO2 plume in the desired way. This ability to manipulate the CO2by imposing pressure gradients has the potential to substantiallyreduce the inherent risk of geologic sequestration of CO2 (Buschecket al., 2011).
The permeate (low salinity water) produced from the reverseosmosis desalination process would be suitable for power plant
cooling, or could be used for agricultural, domestic, or other usesfor which it has monetary value. The residual brine produced bythe process could be re-injected back into the aquifer, such that
1320 W.L. Bourcier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1319–1328
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ig. 1. Schematic of reverse osmosis membrane geometry. Example pressures of feedreatment of seawater.
here would be minimal brine disposal cost. With brine re-injection,he net volume of water removed would be approximately equalo the amount of fresh water recovered in the reverse osmosisrocess.
In this paper, we examine the costs and benefits associated withhe concept of removal and desalination of fluids present in CO2equestration sites. We note that the full development of all ofhe concepts noted in this paper will require a considerable bodyf future work. The scope of this paper is necessarily limited tonly one aspect, water treatment. Using brine removal to con-rol reservoir pressure and manage plume shape and migration iseing developed within the concept of Active CO2 Reservoir Man-gement, or ACRM (Buscheck et al., 2011; Court et al., 2011). Theeed to manage displaced brine has been noted in the case of com-ercial scale sequestration in the Rock Springs Uplift of Wyoming
Surdam et al., 2011). Court et al. (2011) note other examples wherehis need has been realized. We note that the issue of pressureuildup in saline formations targeted for CO2 sequestration hasecome a somewhat controversial topic (cf. Ehlig-Economides andconomides, 2010; Dooley and Davidson, 2010), and that morend specific analyses will be required to determine where pressureuildup (or brine migration, which relieves pressure buildup) mightnd might not be a potential problem and under which conditionst may be possible to produce pressurized brine.
.1. Desalination technologies
For most of human history, desalination of water has beenccomplished through distillation. Most of the world’s desalinationf ocean water is still carried out using thermal distillation (Buros,000). However, membrane-based methods such as reverse osmo-is, developed over the last 50 years, are now the technology ofhoice for new desalination plants (NRC, 2008). Membrane-basedethods are expected to soon overtake thermal methods in total
apacity, driven mainly by the high energy costs of thermal meth-ds (Karagiannis and Soldatos, 2008). The heat of vaporizationf water is so large (∼630 kW-h/m3) that even extremely effi-ient energy recovery (>90%) in thermal systems cannot competeith energy costs of 4–6 kW-h/m3 typical for seawater RO systems
Khawaji et al., 2008).Reverse osmosis (RO) uses a membrane permeable to water
ut impermeable to salt to separate salt from water. To carry out
everse osmosis, a pressure gradient is imposed across a membranehat separates the solution to be de-salted (e.g. ocean water) frommore dilute solution. Because of the osmotic pressure gradient,ater naturally tends to flow from the dilute solution to the saltentrate and permeate are shown to illustrate relative pressures during conventional
solution. By applying a pressure larger than the osmotic pressurein an opposite direction, water can be forced to flow from the saltsolution across the membrane to the dilute side. For desalinationof seawater, pressures on the order of 5–7 MPa (700–1200 psi) areused to overcome the 2.5 MPa (∼360 psi) of osmotic pressure gra-dient between ocean water and dilute permeate (Fritzmann et al.,2007).
Typical aromatic polyamide thin film composite reverse osmosismembranes are composed of a 200 nm layer of semi-permeablepolyamide that is cast onto a porous polysulfone support material(Fig. 1) which itself sits on a spacer to allow water flow along themembrane. The pore size of the polyamide is controlled for thespecific application of use. Pores size estimates are in the rangeof 5–10 A (Baker, 2004). Smaller pore-sized (tighter) membranesare used for high salinity solutions such as seawater in order toincrease salt rejection. Larger pore-size membranes are used formore dilute (brackish) waters where less salt rejection is needed toproduce permeates that are potable (<500 mg/L TDS). Salt rejectionfor tight membranes is greater than 99%; salt rejection for brackishwater membranes is 98–99% (Amjad, 1993). The downside to highrejection is that the membrane is less permeable for water transportand therefore higher driving pressures are needed to produce thesame water flux (Fritzmann et al., 2007).
Reverse osmosis is limited by the maximum pressure differ-ence that the membrane spacers can withstand without collapsingand causing the membrane to become impermeable to waterflow (Matsuura, 2001). The current limit for commercially avail-able membranes is about 8.3 MPa (1200 psi). The semi-permeablepolyamide layer can withstand higher pressures and does not itselfcurrently limit maximum RO pressures.
Nanofiltration (NF) membranes are similar to RO membranes,but are synthesized with larger pores and usually bear some surfacecharge due to changes in the membrane chemistry. They are able toseparate divalent from monovalent ions. For example, a nanofiltra-tion membrane may allow Na and Cl to pass through the membranefreely, while rejecting greater than 90% of the sulfate and calciumin many fluids (Hilal et al., 2004). The actual selectivity of NF mem-branes varies, with some functioning more like RO membranceswith low rejection. NF membranes are typically used for removinghardness (divalent cations such as calcium) from waters, and alsofor removing sulfate to prevent sulfate mineral scaling. Because NFremoves only some fraction of the electrolyte content, it operates
at lower pressure differences than RO (because the osmotic pres-sure depends on the salinity gradient across the membrane, not thetotal salinity). NF can be used in front of RO to partially reduce thefeed salinity, and reduce pressure for the subsequent RO stage. In
of Greenhouse Gas Control 5 (2011) 1319–1328 1321
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W.L. Bourcier et al. / International Journal
his way, the maximum salinity limit for membrane treatment cane increased.
Reverse osmosis systems can operate over a range of waterecoveries. Recovery is defined as the volume ratio of product waterpermeate) to feed. High recoveries are preferred because they
ore efficiently use the energy applied to pressurize the fluid. Onceeed water is pressurized, it can continue to generate a clean per-
eate until its salinity increases to the point where the osmoticressure gradient is nearly equal to the applied pressure. Upon exit-
ng an RO unit, the residual brine (on the feed side of the membrane)s still pressurized. Energy recovery methods are generally used toapture that energy, using either turbines, or pressure exchangershat transfer the pressure of the exiting concentrate to the upstreameed.
The downside to high recoveries is the increased likelihoodf mineral scaling. High pressures cause more rapid water fluxhrough the membrane and increases “concentration polariza-ion” at the membrane surface. This refers to the buildup at the
embrane surface of anything contained in the fluid that doesot pass through the membrane, in particular salt and high MWpecies too large to fit through the 5–10 A pores in the polyamideayer.
All of these considerations must be accounted for in access-ng desalination costs. Software packages such as WTCOST (Moch,008) have been created that provide for unit process costsf each treatment step and are used to identify the minimumost treatment configuration. Seawater desalination costs forembrane-based systems generally vary between $0.50 and $1.00
er m3 of product (Karagiannis and Soldatos, 2008). Costs increaseith the salinity of the fluid because of the higher pressureeeded to drive water through the membranes. Energy costs foreawater desalination are generally half the total cost. Becauseembranes are easily fouled by organics, biological activity, andineral formation, pretreatment to avoid fouling is needed and
an significantly increase cost. Pretreatment is highly specific to theeed, with waters having high organic productivity being particu-arly difficult to pretreat. Formation waters are generally not veryiologically active and some of these pretreatment costs can there-ore be avoided. However, formation waters may contain otheromponents that would introduce a different set of pretreatmentroblems and costs. We will discuss this later in this paper.
.2. Limitations of membrane desalination technologies
The main limitation of membrane-based desalination technolo-ies is the salinity of the feed. Higher salinity requires a higherressure difference across the membrane in order to produce per-eate. This pressure difference must be maintained below a certain
imit to avoid loss of permeability due to membrane compaction.s noted previously, this limit is associated with the properties of
he spacers. Fig. 2 shows the relationships between osmotic pres-ure and water recovery (percent extraction) during treatment ofaters having a range of starting salinities. A recovery of 20% means
hat for each 1000 mL of feed, 200 mL of permeate and 800 mL ofrine are produced. Lines are drawn for recoveries of 0, 10, 20, and0%. This plot illustrates the basic effect, which is driven mainly byalinity as represented by TDS (Total Dissolved Salts), commonlyiven in mg/L. All the waters here are taken to be in the “seawa-er” family (which extends in nature to about 350,000 mg/L TDS,reater than the range shown here). The ionic ratios correspondo those in seawater, the essential feature being that these brinesre NaCl-dominated. The results here are representative for many
ut not all subsurface brines (cf. Kharaka and Hanor, 2007). Ateast two other major types occur: Na–Ca–Cl brines and Na–Cl–SO4rines, which occur in the Rocky Mountain Basins region. Some-hat different results should be expected for waters of these types
Fig. 2. Osmotic pressure for salt solutions as a function of system recovery.
(we will present below examples of calculations for specific brinechemistries).
Fig. 2 can be used to estimate the maximum osmotic pressurefor a fluid undergoing desalination at a given recovery (it will bemore accurate for a fluid in the seawater family). For example, for asalinity of 78,000 mg/L (shown by the left vertical dotted line), theinitial osmotic pressure is about 6.2 MPa (900 psi), and after recov-ering 20% of the fluid as permeate, the residual brine would havean osmotic pressure of 8.2 MPa (1200 psi), the upper maximum forcurrent RO membranes. Above this pressure the membrane prod-uct spacers collapse and the membrane elements can no longer passpermeate to the discharge tube (see Fig. 1).
Note that this pressure limit is not a limit for the functionalpart of the membrane, only the materials currently used as sup-ports. Higher pressure membranes useful up to 10.3 MPa (1500 psi)have been developed and allow brines up to about 80,000 TDSto be desalinated (e.g. Baker, 2004) but are not yet commerciallyavailable.
2. Subsurface brines
To estimate brine desalination costs, we need to know the char-acteristics of the fluids present in the subsurface at proposed carbonstorage sites. Because a catalog of site data are not yet available,we use here a database for subsurface brines co-produced from oiland gas wells and compiled by the U.S. Geological Survey (Breit,2002). This dataset covers the oil and gas-producing areas of theU.S. and has over 14,000 analyses. These data are relevant becausethese brines are typical of sedimentary basins that host petroleumreservoirs in permeable rocks, sites that are also targets for subsur-face CO2 storage (Kharaka and Hanor, 2007). Analysis of these dataprovides compositional information on the brines present in sedi-mentary basins, and their distribution with depth and geographiclocation. The U.S.G.S. database contains abundant data for most ofthe states west of the Mississippi River, but does not have good cov-erage for most of the eastern states. We note that other databasesof produced waters exist, notably NATCARB (Carr et al., 2007).
Fig. 3 shows the distribution of salinities from the U.S.G.S.database for fluids produced from depths greater than 3000 feet(about one kilometer). The salinity range varies widely up to over300,000 mg/L (30 wt%) where salt mineral saturation begins to limitTDS. Over half of the brines have salinities less than 85,000 mg/LTDS. Although there is some tendency for brines to become moreconcentrated with depth, there are many exceptions to this. For
example in Wyoming, waters in the Nugget Formation with salini-ties of over 80,000 mg/L TDS overlie the Tensleep Formation, whichhas waters with salinities less than 25,000 mg/L TDS (cf. data forSublette County, Wyoming, from Breit, 2002).
1322 W.L. Bourcier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1319–1328
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ig. 3. Histogram of salinities of produced waters in the U.S. from depths greaterhan 3000 feet. Data from the U.S.G.S. database of produced waters (Breit, 2002).
The concentration of calcium plus magnesium (hardness, cal-ium + magnesium, expressed as equivalent mg/L calcium) in theatabase is shown in Fig. 4. Fluids with relatively high hardnessre more susceptible to scale formation, but are also amenable toalcium and magnesium removal using nanofiltation. Fluids withardness above about 10,000 mg/L are prime candidates for salin-
ty reduction using an NF-RO process which effectively increaseshe salinity maximum for membrane desalination by first remov-ng hardness (calcium, magnesium, and other divalent ions) prioro removal of the remaining salts. Over 20% of produced watersave hardness in this range based on the U.S.G.S. database.
Some produced brines have significant hydrocarbon contents,hich will require complete removal as these compounds are
nown to irreversibly foul RO membranes (Fritzmann et al., 2007).his is one of the reasons many co-produced waters from fossil fuelroduction are not currently treated, but are instead re-injected
nto the subsurface, at significant cost. For this reason, fluids fromO2 sequestration sites that contain significant amounts of hydro-arbons will be more costly to desalinate. Additional research iseeded in this area to address this problem. New treatment meth-ds may be needed if existing ones prove insufficient or overlyostly. Other pre-treatment needs may arise from the presence ofther components, such as dissolved iron and silica. In the presentaper, we make the assumption that such problems can be dealtith at reasonable cost, and focus instead on mineral precipita-
ion involving major ions and the development of osmotic pressureesulting from extraction of water.
Both mineral scaling and osmotic pressure can limit how muchresh water can be recovered from a given brine. An example is
hown in Figs. 5 and 6 for reverse osmosis treatment of threeolutions: seawater, and two saline waters from formations inublette County, Wyoming, a potential carbon storage site (cf.ig. 4. Histogram of hardness (Ca + Mg expressed as mg/L Ca) for produced watersn the U.S. using data from the U.S.G.S. database of produced waters (Breit, 2002).
0.0
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Glauberite
Nahcolite
Thenardite
Fig. 5. Predicted mineral precipitation as a function of water removal due to reverseosmosis at 50 ◦C in a batch system for (a) seawater brine (35,000 mg/L); (b) SubletteCounty Wyoming #3 brine (86,000 mg/L); and (c) Sublette County Wyoming #2brine (24,500 mg/L). Minor calcite (CaCO3) precipitates at the start (not shown) andpersists for all three fluids.
W.L. Bourcier et al. / International Journal of Gre
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Fig. 6. Predicted osmotic pressure in residual brine as a function of water removaldue to reverse osmosis at 50 ◦C in a batch system. (a) Seawater; (b) SubletteCounty Wyoming #3 brine (86,000 mg/L); (c) Sublette County Wyoming #2 brine(24,500 mg/L). The “Pitzer” curves show the osmotic pressure as obtained fromthe activity of water calculated from a thermodynamic model using Pitzer’s equa-tions incorporated in the EQ3/6 modeling code. The Dow model uses the van’t Hoffequation discussed in text.
enhouse Gas Control 5 (2011) 1319–1328 1323
http://www.electricityforum.com/news/nov08/DOEtofundcarbo-ncaptureproject.html). The compositions of the two saline forma-tion waters were taken from Breit (2002). The first Wyoming brine,“Sublette #3” (UNIQID 49003037), from the Nugget Formation,is very saline (86,000 mg/L TDS, 29,000 Na+, 2650 Ca++, 2400 K+,1200 SO4
−, 415 HCO3−). The second Wyoming brine, “Sublette #2”
(UNIQID49003313), is a high sulfate brine of lower total salinity(24,500 mg/L TDS, 7800 Na+, 200 Ca++, 100 Mg++, 14,600 SO4
−,1340 HCO3
−) from the Tensleep Formation, which underlies theNugget Formation. In Fig. 5 we show predicted mineral scaling ofseawater and the Sublette County brines during reverse osmosistreatment simulated using the EQ3/6 modeling software (Woleryand Jarek, 2003). For seawater, Fig. 5a shows that as the wateris concentrated during reverse osmosis, both calcite (CaCO3) andanhydrite (CaSO4) are predicted to precipitate. However, bothphases can be controlled using anti-scalants and do not necessarilylimit water recovery. Likewise for the Sublette County #3 brine,calcite and anhydrite are predicted to precipitate (Fig. 5b). For thehigh sulfate Sublette #2 brine (Fig. 5c), only anhydrite is predictedto precipitate even at recoveries of over 85%. However, in all threecases, at very high recoveries, halite (NaCl) and other phases suchas glauberite (Na2Ca(SO4)2), nahcolite (NaHCO3), and thenardite(Na2SO4) are predicted to precipitate in quantitites that provide afirm upper limit to recovery.
We calculate the osmotic pressure for the same feed solutionsas they are concentrated in the reverse osmosis process. The pres-sures are shown by the red and blue lines in Fig. 6. The blue line iscalculated using a simple formula based on ideal mixing (the van’tHoff equation; Dow, 2009). The red line shows a more rigorous cal-culation using Pitzer’s equations and is based on thermodynamicmeasurements of salt solutions and includes non-ideal mixing cor-rections (Greenberg and Møller, 1989; Pabalan and Pitzer, 1987).As can be seen, at high salinities, the two methods differ by severalMPa, indicating that the more rigorous osmotic pressure calcula-tion must be used for solutions for modeling salinities above about1 molal. Comparing Fig. 6b and c shows that the more rigorous cal-culated pressure can be either higher or lower than the van’t Hoffvalue depending on the solution composition, and in particular onthe relative amounts of monovalent and divalent ions.
The upper limit on recovery occurs when the lines cross thehorizontal line drawn at 8.3 MPa (1200 psi), the current membranetechnology pressure limit. For seawater (a), the absolute limit is arecovery of about 65%, which corresponds to a practical recoverylimit of less than about 50%. This is because for normal operation,the pressure across the membrane must be equal to the osmoticpressure plus an additional differential pressure of 0.7–1.4 MPa(100–200 psi) to maintain appreciable flow of water across themembrane. For the more saline Sublette #3 water (∼86,000 mg/LTDS) the theoretical recovery is limited by osmotic pressure toabout 7%, or for practical purposes, the brine is at the edge oftreatability using conventional reverse osmosis. For the less salineSublette #2 water, recoveries of 80% appear achievable. Overall,it is apparent that the development of osmotic pressure is morelimiting than mineral precipitation.
We conclude that the potential for freshwater extraction fromsaline brines can be categorized as follows:
• 10,000–40,000 mg/L TDS: standard RO with ≥50% recovery;• 40,000–85,000 mg/L TDS: standard RO with ≥10% recovery;
higher recovery possible using 10.3 MPa (1500 psi) RO mem-branes and/or multi-stage incremental desalination including NF(nanofiltration);
• 85,000–300,000 mg/L TDS: multi-stage process (NF + RO) usingprocess design that may differ significantly from seawater sys-tems;
• >300,000 mg/L TDS brines: not likely to be treatable.
1324 W.L. Bourcier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1319–1328
seaw
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Fig. 7. Comparison of major treatment steps in conventional high-pressure
As can be seen in Figs. 3 and 4, brines in the 10,000–85,000 mg/LDS range appear to be abundant (geographically and with depth)nd could be targeted in siting CCS operations. These brines wille chemically different than seawater and these differences willffect treatment costs. Brines greater in salinity than 85,000 will bexpensive to treat and will not be addressed further in this paper.
.1. Desalination of seawater vs. saline formation water
Because many of the candidate brines present in the subsur-ace at sequestration sites have salinities somewhat similar to thatf seawater, we will use for our first cost analysis a generic fluidaving a total salinity equal to that of normal seawater, but withpecific chemical characteristics more typical of subsurface brines.his allows the most accurate cost estimate to be made with min-mal additional assumptions. There is abundant experience andost data available for the reverse osmosis desalination of seawaterKaragiannis and Soldatos, 2008).
A conventional seawater RO desalination system uses raw oceanater as feed. The ocean water is first filtered for particulates
nd micro-organisms. Instead of open water intakes, many treat-ent plants now take the feed from wells drilled into the coastal
and wedge to take advantage of the sand as a first filtration stepFrenkel, 2011). The water may then be acidified, have antiscalantsdded, filtered again at a finer size, and then pressurized and passednto the RO elements. The RO elements are arranged in multipletages to maximize permeate recovery and minimize the energysed to pressurize the system. The feed water stays pressurizedhroughout the desalination process. It exits with a pressure only–30 kPa lower than the starting pressure. The permeate fluid maye adjusted in pH and chlorinated prior to being sent to storage andistribution. The concentrate is sent to disposal, typically piped outo sea.
The major steps in desalination of saline formation water areompared with those of conventional seawater RO in Fig. 7. Forhe saline water, we assume for this analysis that we have a well-
ead pressure of 8.3 MPa (1200 psi). This is at the upper limit ofhat can be used in current RO systems and therefore providesn end-member cost comparison. We also account for the likelyifferences for desalinating fluids pumped from the subsurface vs.
ater desalination vs. low-recovery desalination of saline formation waters.
those obtained from the ocean. We then compare the process stepsfor seawater vs. pressurized brine and examine how they differ.
The first difference is that the saline formation fluids areassumed to be saturated with several mineral phases, a goodassumption since they have had significant time in the subsurfaceto equilibrate with the surrounding reservoir rocks. These miner-als include carbonates, sulfates, and silicates. As a consequence weinclude in our process costs much more extensive pre-treatment,in particular the addition of antiscalants and acid, to avoid mineralscaling of the membranes.
We assume there is none or only minor biological activity in thesaline formation fluids, as is typical for fluids co-produced with oiland gas compared with ocean water. This allows us to eliminatethe need for extensive biological pre-treatment to avoid biofilmformation. Biofilm formation is generally regarded one of the mostdifficult problems in seawater RO systems (Amjad, 1993).
Many saline formation waters contain significant hydrocarbons,which are known to aggressively foul RO membranes (Fritzmannet al., 2007). Removal methods may be needed if the saline watercontains significant hydrocarbons.
Since the measured temperature for many reservoir feeds canexceed 100 ◦C, the saline formation water may need to be chilled toa temperature cool enough to allow treatment using conventionalRO membranes. The current limit for long-term use of conventionalmembranes is around 50 ◦C, although high-temperature elementsare available that can be used up to 70 ◦C (Snow et al., 1996). Heatexchangers will therefore be necessary to bring the feed water towithin the membrane operating envelope. We include in our coststhe use of heat exchangers of a shell and tube design (due to thehigh pressures involved) and made of 317L stainless steel or betterand employed in a split stream formation. The necessary corro-sion resistance of the material to be used will be a function of thechloride and sulfide levels in the reservoir fluid.
The split heat exchanger arrangement would cool a portion ofthe feed using the counter flowing, previously cooled permeate.The second heat exchanger package would treat the remaining hot
feed by cooling it with the previously cooled reject.The inherent heat losses and inefficiencies of this heat trans-fer arrangement would be supplemented by an auxiliary heattransfer device, which would either be in the form of an air cooled
W.L. Bourcier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1319–1328 1325
Table 1Conventional seawater vs. saline formation water reverse osmosis treatment systems.
RO system major component/items Conventional SW systems Reservoir treatment system
EquipmentHigh pressure feed piping Only necessary after the high pressure pump Starting from the well headHigh pressure reject piping Only to back pressure control valve, i.e. skid boundary Return line to the deep well injection pointIntake/feed structure Open intake structure or sea wells Reservoir wellsPretreatment system Multi-media filter or micro-filter Activated carbon filterFeed pumping system Needed from feed source to RO skid Not needed – use well head pressurePretreatment system Cartridge filters Cartridge filtersHeat exchanger None Yes. Reduce feed temperature to 50 ◦CChemical Injection system – bio control Yes – for bio control followed by reducing agent Not neededChemical Injection system – scale control Yes – scale inhibitor for calcium carbonate Yes – scale inhibitor for calcium carbonate and/or calcium
sulfateHigh pressure pump Yes (most costly item of RO skid) No - Use well head pressureEnergy recovery device Yes – typically used for larger systems No – keep reject at maximum pressure for reinjectionRO pressure vessels (PV) Same for equal permeate capacity Same for equal permeate capacityRO membranes Conventional seawater membranes Conventional seawater membranesArray – configuration Single pass 7 elements per PV Single pass 7 elements per PVPercent recovery Ranges between 30 and 50%, depending on seawater
concentration, system size and energy costsRanges between 15 and 45%, depending on reservoir feedconcentration and composition
System I&C Similar to reservoir system Similar to seawater systemMCC/VFD Includes electrical supply for high HP pump(s) Same as conventional SW system, with exception of the
high HP pump(s), i.e. no VFD neededValves, controls and instrumentation Similar to reservoir system Similar to seawater system, except high pressure pump
associated valving and controlsTanking Similar to reservoir system Similar to seawater systemDeep well injection Reject typically goes to seawater outfall Deep well injection requiredCivil works Very site dependent Very site dependentOperationsOperating Similar to reservoir system except no deep well
injectionSimilar to seawater system except no high pressurepump(s) and energy recovery, but use of deep wellinjection for reject concentrate
Energy consumption Energy to deliver feed to high pressure pumps andthen raise to ∼1000 psi
Similar to seawater system without the feed and processpumping. Deep well injection pumping energy required
Chemicals consumption, including membrane Similar to reservoir system Similar to seawater system
acosWse
datip
sbccerblhtCa
teTwt
cleaningMaintenance as % of capital Similar to reservoir system
tmospheric exchanger or it would be tied into the power plantooling system to provide for the additional cooling as needed. Inur cost analysis we assume air cooling. For the latter arrangement,urplus cooling capacity from the power plant would be required.
hile there would be some pressure losses associated with the pas-age through the heat exchange system, the pressure drops acrossxchangers are typically less than 100 kPa.
One area of difference between the two processes is that theesign of seawater systems views the amount of feed strictly fromn economic point of view. The only treatment product of value ishe extracted permeate water. Seawater RO plants strive to max-mize recovery in order to reduce the cost per unit volume ofermeate for a given degree of pressurization.
The situation is different for conditions where there is significantubsurface pressure (e.g. well confined sequestration reservoirs)ecause a significant fraction of the energy needed to drive ROan be obtained from the pressurized reservoir. In this case, energyosts, which ordinarily make it most effective to maximize recov-ry, are less important and allow favorable economics for lowecovery. The benefit from low recovery is that it increases mem-rane longevity by reducing the risk of membrane scaling and
owering the driving pressure. The down side is that one now thenas to pump and treat a larger volume of brine in order to producehe same amount of permeate – and equivalent storage space forO2 in the subsurface. We explored a range of recoveries in our costnalysis.
The last stage of seawater RO is an energy recovery step, wherehe pressure of the exiting concentrate is recovered and used to
ither generate electrical energy or to pressurize the incoming feed.he energy recovery device is not needed for our saline formationater process because the residual pressure is needed to re-injecthe concentrate back into the formation. In fact, the pressure of the
Similar to seawater system
concentrate may need to be increased, depending on where it isinjected into the formation.
3. Saline formation water desalination costs
A list of the treatment components for seawater compared tosaline formation waters is given in Table 1. These are the unit pro-cesses and operations that are considered in our cost analysis thatfollows. The system capital costs include the following major com-ponents:
1. High pressure piping from the well head to the treatment sys-tem. The piping would require a material quality of 316L orbetter to minimize corrosion.
2. Activated carbon pretreatment to remove organics (oil andgrease), unlike conventional seawater RO, these systems willhave to accommodate high feed pressure.
3. Heat exchanger to reduce the feed temperature to 50 ◦C.4. Cartridge filter system.5. Chemical injection system for scale inhibitor or acid.6. RO skids, pressure vessels, and membranes.7. System instrumentation and control (I&C).8. System piping valves, controls & instruments.9. Forward flush and clean-in-place systems.
10. Tankage for chemicals and permeate water.11. Re-pumping station for reject (concentrate) deep well injec-
tion.12. Pretreatment system for deep well injection.13. Civil works, equipment housing and installation costs.14. One spare RO train for standby assurance.
1326 W.L. Bourcier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1319–1328
Table 2Major costs and cost factors for RO treatment systems.
Assumed discount rate 7.00%Plant lifetime/evaluation period 25.00 yearsOperating cost escalation, %/year 0.00%
Base operating pressure at reference conditions 1200 psiBase recovery 40%Cleaning, times/year 0.5Chemicals $3
ConsumablesActivated carbon $30 $/ft3
Filter cartridges $4 $/eachPower, repressurization $0.07 $/kWhPower, misc uses 200.00 kWh/day
LaborFringe and overhead multiplier 2.00 factorSupervisor $150,000 SalaryOperators $60,000 SalaryTechnicians $75,000 Salary
op
123
4
5
tfteep
rAfatsaocwep
ipw$dac
ni
$-$100$200$300$400$500$600$700$800$900
$1,000
50%45%40%35%30%25%20%15%10%
LCO
W -
v $
/AC-
FT
System RO Recovery
Levelized Cost of Water vs. Recovery
2 MGD - LEVELIZED COST/AC-FT, 25 yrs4 MGD - LEVELIZED COST/AC-FT, 25 yrs6 MGD - LEVELIZED COST/AC-FT, 25 yrs
One of the items to warrant further review is the problem ofoptimizing the front end heat exchanger system used to reduce thereservoir fluid’s temperature for RO processing. As described above,
$500
$700
$900
$1,100
$1,300
$1,500
$1,700
50%45%40%35%30%25%20%15%10%System RO Recovery
LCO
W -
$/A
C-F
T
2 MGD - Pump to Pressure 2 MGD - Reservoir Pressure
Mechanical maintenance, % of capital 3%Length of high pressure piping 125 ft
We did not include costs for land acquisition or for constructionf the brine well field, which includes costs associated with drillingroduction and reinjection wells.
Operating and maintenance costs include:
. Consumable chemicals, filter replacements, etc.
. Hazardous wastes disposal costs.
. RO membrane replacement; depreciation ranging over 2–4years, depending on the percent recovery of the system.
. Labor costs for supervisor/engineer, operators and technician.Salaries estimated with an overhead factor of two (2).
. Maintenance estimated at 3% of capital annually.
The capital and operating costs were analyzed for several sys-em sizes and recoveries. Base models for costing were developedor 2 and 4 million gallons per day (MGD) systems, which werehen used to estimate the costs for the 6 and 8 MGD systems, usingxponential escalation factors. Elements, vessels, and piping werestimated by actual cost and quantity or in the case of piping peround costs for calculated pipe diameters and thickness.
RO costs were estimated for 45%, 40%, 35%, 30%, 25% and 15%ecoveries – defined as percent permeate extracted from the feed.s described above, percent recovery is a function of the incoming
eed salinity. Since processing ceiling is dictated by the maximumllowable membrane operating pressure, the osmotic pressure ofhe reject stream becomes the limiting factor. The higher the feedalinity the lower the percent recovery before the maximum allow-ble pressure is reached. Operating pressure is comprised of thesmotic pressure of the concentrate plus the hydraulic pressure lossreated by the membrane barrier, plus a driving pressure to pushater through the membrane. At a feed TDS of 60,000 mg/L, recov-
ry may be limited to 15–20% recovery for maximum operatingressures of 8.3 MPa (1200 psi).
Table 2 summarizes the major costs and cost factors employedn developing the system costs. Results for the scenarios are dis-layed in Table 3 and Figs. 8 and 9. Fig. 8 shows that levelizedater costs per acre-foot of permeate vary from about $450 to
1000 (32–81¢ m3, one acre foot is 1233 m3), and that the costsecrease as the size of the operation increases. In addition, there iscost effective recovery point (minimum) in plots of recovery vs.
ost.Most importantly, however, the lower recovery options are eco-
omically close to the higher recovery options – that is, the curves fairly flat. With conventional pumped systems the curve is much
Fig. 8. Comparison of estimated levelized costs per acre-foot of brine at 4 differentproduction rates in million gallons permeate per day (MGD).
more pronounced, greatly favoring higher recovery. As mentionedabove, this is because the additional costs associated with movingand pressurizing additional feed is now reduced because the pres-surization is no longer a large operating expense. In this scenariowe assume that the energy to pressurize the fluid is supplied bythe reservoir (which we believe may be obtainable for some butnot likely all storage reservoir scenarios).
Although we have assumed a fluid salinity similar to that of sea-water, our estimated costs are not highly sensitive to salinity. Thisis because we assume the energy for desalination comes mainlyfrom the existing reservoir pressure field and is not a major com-ponent of the operating cost. In a conventional desalination system,an increase in salinity makes the process more expensive due to theneed for additional energy to operate the system at higher pres-sures. For this reason, we believe our estimated costs are fairlyrepresentative for the range of fluid salinities for which an essen-tially conventional RO process can be used (i.e. 10,000–85,000 mg/LTDS).
Fig. 9 directly compares the costs for a system using reservoirpressure to drive the RO process with an identical system usinghigh-pressure pumps. The difference between the curves reflectsthe estimated cost of energy to drive the desalination system.Note also the pumped system shows a steeper cost vs. recoverycurve.
Fig. 9. Comparison of levelized cost of water desalination for pressurized brinesfrom CCS sites (where we assume well-head pressure is adequate to drive RO sys-tem) vs. cost for conventional seawater RO that includes energy needed for highpressure pumps.
W.L. Bourcier et al. / International Journal of Greenhouse Gas Control 5 (2011) 1319–1328 1327
Table 3Itemization of costs for water desalination plants of 2 and 8 MGD permeate at 40% water recovery.
Base FEED flow, MGD 5.00 Product flow, MGD 2.0 2.0 2.0 2.0 2.0 2.0Base PRODUCT flow, MGD 2.00 Recovery, percent 45.0% 40.0% 35.0% 30.0% 25.0% 15.0%Base recovery 40% AC-FT/year 2240 2240 2240 2240 2240 2240
Total direct operating costs $/year $1,233,698 $1,206,483 $1,189,693 $1,187,315 $1,183,468 $1,251,878$/AC-FT $550 $538 $531 $530 $528 $558
Installed capital cost, $/GPD $3 $3 $3 $3 $3 $4Total direct capital $3,844,576 $3,912,316 $4,027,492 $4,234,852 $4,588,768 $5,823,751Installed capital cost $6,934,334 $6,987,133 $7,120,062 $7,414,175 $7,962,181 $9,869,483NPV of O&M $14,149,698 $13,837,562 $13,644,994 $13,617,716 $13,573,596 $14,358,217
NPV of O&M + capital cost $21,084,032 $20,824,694 $20,765,056 $21,031,891 $21,535,777 $24,227,701NPV of AC-FT/year NPV AC-FT/period 25,691 25,691 25,691 25,691 25,691 25,691
2 MGD – levelized cost/AC-FT, 25 years $821 $811 $808 $819 $838 $943
Base FEED flow, MGD 20.00 Product flow, MGD 8.0 8.0 8.0 8.0 8.0 8.0Base PRODUCT flow, MGD 8.00 Recovery, percent 46.0% 40.0% 35.0% 30.0% 25.0% 15.0%Base recovery 40% AC-FT/year 8960 8960 8960 8960 8960 8960
Total direct operating costs $/year $2,259,846 $2,142,984 $2,073,987 $2,061,919 $2,041,721 $2,291,635$/AC-FT $252 $239 $231 $230 $227 $255
Installed capital cost, $/GPD $2 $2 $2 $2 $2 $3Total direct capital $10,609,856 $10,796,811 $10,996,812 $11,530,415 $12,499,362 $16,138,799Installed capital cost $20,594,678 $20,756,123 $20,947,918 $21,757,741 $23,349,555 $29,354,228NPV of O&M $25,918,931 $24,578,608 $23,787,260 $23,648,839 $23,417,187 $26,283,538
$45102$44
acbnewtotu
4
sh(rfpewminw
almrRwmtpmasa
NPV of O&M + capital cost $46,513,609NPV of AC-FT/year NPV AC-FT/period 102,763
8 MGD – levelized cost/AC-FT, 25 years $453
cooling device to accomplish the necessary heat reduction mightonsist of a split heat exchanger arrangement, which will have toe augmented via a secondary heat reduction source. The specificature of this supplementary heat rejection will have to be consid-red. The technical and cost impact of coupling the heat reductionith the power plant’s cooling system vs. a stand-alone system
hat uses atmospheric cooling, possibly augmented by spray evap-ration, must be made. Since this would consume RO permeate,hereby reducing the amount of recovered water for power plantse, it may be a significant factor to consider.
. Conclusions
Our analysis indicates that RO plants for CCS brines that areimilar in salinity to seawater can be built and operated for aboutalf of the cost of seawater desalination if sufficient over-pressurefrom CO2 injection) exists to supply the pressure needed to driveeverse osmosis. Over 40% of subsurface brines co-produced withossil fuels have salinities at or below seawater. For a fresh waterroduction of 6 MGD (about 23,000 m3 or about 18 acre feet), westimate costs of about 32–40¢/m3 permeate produced. Withoutell-head energy recovery, the costs are from 60 to 80¢/m3 per-eate, similar to conventional seawater desalination. This analysis
ncludes all surface facilities, transfer pumps, and piping, but doesot consider the cost of the brine extraction and reinjection wells,hich will be site dependent.
The costs for membrane desalination of sedimentary brines thatre higher in salinity than our upper limit of 85,000 mg/L will bearger. It is currently impossible to estimate these costs because
embrane treatment technologies that operate in that salinityange do not exist. However, we believe it is likely that improvedO technologies capable of desalinating more concentrated brinesould quickly become available if sufficient demand existed. Asentioned above, both the pressure and temperature limits are due
o support materials and are not intrinsic limitations of the semi-ermeable polyamide membrane. High temperature-pressure RO
embranes could be constructed with improved support materi-ls. For these reasons, we expect the treatment costs for higheralinity brines to be only modestly higher than those for seawaterfter allowing for higher pressures and lower recoveries.
,334,730 $44,735,177 $45,406,581 $46,766,742 $55,637,766,763 102,763 102,763 102,763 102,7631 $435 $442 $455 $541
The amount of fresh water produced by our process is signif-icant. For example, a typical 1 GW coal plant emits more than7 MtCO2/year. A well-designed capture system might provide6 MtCO2 for sequestration. Sequestered at a depth of 300 m, thisCO2 would displace about 8 million m3 of water per year, or about22,000 m3/day. Reverse osmosis treatment of that brine with arecovery of 40% would annually produce about 3.2 million m3 offresh water, which could serve the needs of 10,000 homes, irrigate2000 acres of cropland, or provide half of the water volume usageof a typical 1 GW IGCC power plant, based on a plant use of about2000 L/MW-h (Grubert and Kitasei, 2010).
An additional benefit of our process is that strategic withdrawaland reinjection of saline aquifer brines will allow manipulation ofthe location and migration of the CO2 plume in the subsurface.This could be accomplished by using brine injection/withdrawalto alter the subsurface pressure field so as to affect the pressuregradient surrounding the CO2 plume in the desired way. This abil-ity to manipulate the CO2 by imposing pressure gradients hasthe potential to substantially reduce the inherent risk of geologicsequestration of CO2.
Finally, site selection for saline aquifers suitable for disposal ofcarbon dioxide and for which brine removal increases reservoircapacity should include some consideration of the fluid salinity,because lower salinity fluids are less costly to desalinate.
Acknowledgements
Financial support for this work was provided by DOE throughthe National Energy Technology Laboratory. We thank Sean Plasyn-ski and Andrea McNemar for their help and advice. We also thankGeoffrey Thyne of the Enhanced Oil Recovery Institute at the Uni-versity of Wyoming, George Breit of the USGS, and Julio Friedmannof LLNL for helpful discussions about produced waters.
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