Electrochemical Ion-Exchange Regeneration and Fluidized Bed

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Electrochemical Ion Exchange Regeneration and Fluidized Bed Crystallization for Zero Liquid Discharge Water Softening Yingying Chen, Jake R Davis, Chi Huynh Nguyen, James Claude Baygents, and James Farrell Environ. Sci. Technol., Just Accepted Manuscript • DOI: 10.1021/acs.est.5b05606 • Publication Date (Web): 10 May 2016 Downloaded from http://pubs.acs.org on May 10, 2016

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Environmental Science & Technology

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Electrochemical Ion Exchange Regeneration and Fluidized Bed

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Crystallization for Zero Liquid Discharge Water Softening

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Yingying Chen, Jake R. Davis, Chi H. Nguyen, James C. Baygents and James Farrell*

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Department of Chemical and Environmental Engineering

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University of Arizona, Tucson, AZ 85721

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Keywords: bipolar membrane electrodialysis, ion exchange regeneration, water softening, fluidized bed crystallization, zero liquid discharge, brine disposal, Tokuyama BP1, Fumasep FBM

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*corresponding author: [email protected]; 520 940-0487

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ACS Paragon Plus Environment

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Abstract

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This research investigated the use of an electrochemical system for regenerating ion

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exchange media and for promoting crystallization of hardness minerals in a fluidized bed

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crystallization reactor (FBCR). The closed-loop process eliminates the creation of waste brine

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solutions that are normally produced when regenerating ion exchange media.

34

membrane electrodialysis stack was used to generate acids and bases from 100 mM salt solutions.

35

The acid was used to regenerate weak acid cation (WAC) ion exchange media used for water

36

softening. The base solutions were used to absorb CO2 gas, and to provide a source of alkalinity

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for removing noncarbonate hardness by WAC media operated in H+ form. The base solutions

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were also used to promote crystallization of CaCO3 and Mg(OH)2 in a FBCR. The overall

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process removes hardness ions from the water being softened and replaces them with H+ ions,

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slightly decreasing the pH value of the softened water. The current utilization efficiency for acid

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and base production was ~75% over the operational range of interest, and the energy costs for

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producing acids and bases were an order of magnitude lower than the costs for purchasing acid

43

and base in bulk quantities. Ion balances indicate that the closed-loop system will accumulate

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SO42-, Cl- and alkali metal ions. Acid and base balances indicate that for a typical water, small

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amounts of base will be accumulated.

A bipolar

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Introduction

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Many industrial and domestic uses of water require water softening in order to avoid scale

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formation on heat transfer surfaces. Examples of this include, water used in boilers for electric

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power generation and water used in evaporative cooling systems. The most common method of

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industrial and domestic water softening is ion exchange. The regeneration of ion exchange

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media produces large quantities of brine. A single regeneration cycle produces several cubic

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meters of brine solution per cubic meter of ion exchange media.1 Brine disposal methods include

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metered discharge to the sewer system, ocean discharge, evaporation ponds and thermal

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crystallization. 2 Metered discharge to sewer systems can result in toxicity to biomass and

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inhibition of settling in clarifiers used in secondary wastewater treatment.2 Inland areas do not

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have the option of ocean discharge, and thermal crystallization is very expensive. Evaporation

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ponds are also problematic in that they require large land areas and become nuisances due to

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algal growth, attraction of brine flies, and anaerobic microbial growth that results in production

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of a variety of foul smelling and toxic gases (e.g., H2S).

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In most cases, ion exchange brines are sent to the sanitary sewer system for disposal. This

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practice contributes to increasing salinification of public water supplies. The problem of water

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salinification is particularly acute in the desert southwest of the United States. For example, the

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water supplying the Phoenix, Arizona metropolitan area brings in 1.46 million tons of salt each

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year into the local watershed. Only 0.36 million tons of salt flow out of the watershed each

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year.3 This leaves a net of 1.1 million tons of salt accumulation per year, 40% of which ends up

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in the groundwater and 22% in the vadose zone.3 This accumulation does not even include salts

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added to the water as a result of water softening, which is responsible for an estimated 26% of

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the total salt content of wastewater discharged into sanitary sewer systems.3


The high

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contribution of water softening to water salinification has prompted many communities to ban

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water softeners that use salts or acids for regeneration.4

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Bipolar Membrane Electrodialysis

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The problems and costs associated with brine disposal can be eliminated if the acids and

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bases used for ion exchange media regeneration could be reused in a zero liquid discharge,

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closed-loop process. Bipolar membrane electrodialysis (BMED) has been used to produce acids

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and bases from concentrated salt solutions, from reverse osmosis concentrates, and from brine

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solutions resulting from spent acids and bases.5,6,7,8 The repeating unit cell in a BMED stack

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consists of an anion exchange membrane, a bipolar membrane, and a cation exchange membrane,

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as illustrated in Figure S1 in the Supporting Information. BMED stacks can contain up to 300 of

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these unit cells. Bipolar membranes consist of three layers: a strong acid cation exchange layer

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(CEL), a strong base anion exchange layer (AEL), and a very thin intermediate layer containing

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a catalyst. The catalyst layer promotes water dissociation into H+ and OH- ions via the second

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Wien effect.9 Application of a reverse biased electric field across the bipolar membrane results

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in electromigration of H+ ions towards the cathode and OH- ion towards the anode. An anion

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exchange membrane (AEM) traps acid in the acid chamber, as illustrated in Figure S1. Likewise,

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a cation exchange membrane (CEM) traps OH- ions created by water splitting in the base

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chamber. Electromigration of anions from the feed solution through the AEM into the acid

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chambers maintains electroneutrality of the acid.

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through the CEM maintains the charge balance in the base chamber.

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Fluidized Bed Crystallization

Cation migration from the feed solution

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Fluidized bed crystallization reactors (FBCR) have been used to promote crystallization of

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minerals on seed crystals or sand grains suspended in an upflow reactor.10,11,12 In water softening


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applications this process is known as pellet softening, and garnet sand is often used as the seed

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material to promote heterogeneous crystallization. In pellet softening, water containing hardness

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ions are fed into a FBCR along with a stream of Na2CO3, or if the water has sufficient

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bicarbonate alkalinity, a stream of NaOH.

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heterogeneous crystallization of hardness minerals in the reactor, forming pellets. Once the

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crystals grow to ~2-5 mm in diameter, the pellets are flushed out of the bottom of the reactor.

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Due to the large size of the pellets, they are easily dewatered via gravity drainage. Pellet

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softening has been used in the Netherlands for softening at municipal water treatment plants

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since the early 1980s.12 Currently, pellet softening is used at 25 water treatment plants in Europe,

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Asia and Australia, but at only three locations in the United States.13

The seed material and high pH promote

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Two issues have prevented wide adoption of FBCR softening. These are: 1) the need to

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add pH adjusting chemicals, such as NaOH or Na2CO3, and 2) the need for tall reactors, up to 12

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m in height.14 The addition of pH adjusting chemicals is costly and adds to the TDS load of the

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water. Although the softening processes proceeds quite rapidly, with 90% of the hardness

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removal occurring after about 30 seconds of contact time, tall reactors are needed in order to

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ensure sufficient hydraulic detention time that the water exiting the top of the reactor is only

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marginally supersaturated with respect to calcite.11,15 If the saturation index of the water exiting

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the top of the reactor is too high, acid must be added to reduce the degree of supersaturation.

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Acid addition is undesirable because of its cost and the associated increase in treated water

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salinity. For these reasons, acid is not generally added in FBCR softening, and tall reactors are

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employed.10 Even with reactors as tall as 5 m, the level of hardness removal is typically

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~50%.12,14,16


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The goal of this research was to investigate the feasibility of regenerating ion exchange

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media in a closed-loop system that produces no liquid waste products. A BMED stack was used

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to generate acid and base streams using 100 mM Na2SO4 or NaCl solutions. The acid and base

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were used to regenerate ion exchange media. Base generated by the BMED stack was also used

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to produce a high pH stream containing dissolved carbonate ions.

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crystallize CaCO3 and Mg(OH)2 minerals on garnet sand grains. Ion balances and acid/base

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balances indicate that the process is technically feasible. Electrical energy costs for generating

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the acids and bases were low compared to the costs for purchasing acid and base in bulk

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quantities.

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Material and Methods

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Acid and Base Generation

A FBCR was used to

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Acid and base solutions were generated using feed solutions consisting of 100 mM Na2SO4

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or 100 mM NaCl. Solutions were prepared in 18 MΩ cm ultrapure water (UPW) using ACS

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grade reagents obtained from Fisher Scientific.

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continuously stirred 55-gallon high density polyethylene (HDPE) drum. The feed solutions were

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fed into a BMED stack containing 8 unit cells, each consisting of a bipolar membrane, anion

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exchange membrane and cation exchange membrane. The BMED stack was housed in a PC-Cell

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EDQ380 electrodialysis unit (PC-Cell GmbH, Heusweiler, Germany), and had 380 cm2 of active

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area per membrane. Polypropylene mesh spacers with built-in silicone gasketing were placed

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between each membrane pair forming 195 mm x 195 mm x 0.5 mm flow channels. The spacing

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between the electrodes and the cation exchange membranes terminating each end of the stack

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was approximately 2 mm. A 200 mm by 200 mm expanded stainless steel mesh was used as the

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cathode. The anode was the same size and consisted of an expanded titanium mesh with a

The feed solutions were contained in a


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platinum-iridium oxide coating. The AEM (Neosepta ACM) was 0.17 mm thick and had an

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exchange capacity of 1.4-1.7 mequiv g-1. The CEM (Neosepta CMX) was 0.18 mm thick and

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had an exchange capacity of 1.5-1.8 mequiv g-1.

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Fumasep FBM, PEEK reinforced membrane and a Neosepta BM-1.

Two bipolar membranes were tested, a

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Two centrifugal pumps placed in series were used to feed the salt solutions to the BMED

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stack. A volumetric flow rate of 0.5 L min-1 was used for the diluate and flow rates of 0.25 L

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min-1 were used for the acid and base streams. The electrode rinse solution passed through both

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anode and cathode chambers at 1.5 L min-1, and was recirculated to the feed drum after passing

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through a custom degasser and a 10-inch long block carbon filter. The block carbon filter served

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to remove hypochlorite or peroxodisulfate formed via oxidation of Cl- or SO42- at the anode. In

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practice, the electrode wash solutions are normally in a separate loop using a NaOH electrolyte

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that avoids hypochlorite or peroxodisulfate production. The use of the same electrolyte in all cell

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compartments was done to simplify the experimental setup. The BMED cell was operated

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galvanostatically using a Xantrex XRT 40-21 power supply.

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Fluidized Bed Crystallization Reactor

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The FBCR reactor consisted of a 280 cm tall by 6.35 cm diameter clear PVC tubing

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containing a 115 cm long bed of 40-60 mesh garnet sand.

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heterogeneous nucleation sites for hardness mineral crystallization. During the crystallization

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experiments, water was recirculated through the reactor at a flow rate of 0.46 L min-1 using a

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peristaltic pump, as illustrated in Figure 1. This resulted in fluidization of the bed and expansion

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by 15-20%.

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hardness ions or carbonate ions into the base of the reactor, each at flow rate of 0.25 L min-1.

The garnet sand served as

Crystallization experiments were performed by feeding solutions containing


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Solutions exiting the FBCR were first passed through a 6.35 cm diameter by 25 cm long

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Doulton ceramic microfilter contained in standard filter housing. Effluent from the filter was

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passed through the weak acid cation (WAC) ion exchange media in H+ form. This served to

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remove residual hardness ions, lower the pH of the effluent solution to ~5.5, and to convert

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CO32- and HCO3- ions into H2CO3. A second membrane contactor (membrane contactor #2) was

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operated as a liquid-liquid contactor with the effluent from the FBCR passing through the tube

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side. Base solutions with pH values ranging from 11 to 12.5 were passed through the lumen side

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of the contactor. Dissolved CO2 gas from solutions exiting WAC #3 were absorbed by the high

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pH solution. This process was used to reduce the carbonate content of the water that was cycled

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back into feed solution for the BMED stack.

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171 172 173 174 175 176

Figure 1. Schematic diagram of closed-loop system for regenerating ion exchange media and crystallizing hardness minerals in a fluidized bed crystallization reactor (FBCR). A modified version of this figure that contains the composition of each stream is included as Figure S2 in the Supporting Information. Recirculation loops for the BMED cell and FBCR are noted with bold lines.

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Regeneration Experiments


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Weak acid cation ion exchange media (Purolite C104) loaded with hardness ions was

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regenerated at a flow rate of 0.125 L min-1 using acid produced by the BMED stack. The WAC

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media was contained in 6.35 cm diameter by 25 cm long cartridges in standard filter housing.

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Effluent from the WAC cartridges was fed into the FBCR, as illustrated in Figure 1. A lead-lag

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configuration was used to minimize acid addition to the FBCR. WAC#3 can also be regenerated

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in a lead-lag process.

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Base produced by the BMED stack was used to absorb CO2 gas and to regenerate weak

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base anion (WBA) ion exchange media. A fraction of the base solution from the BMED stack

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was passed through the tube side of a Liqui-cel model G420 hollow fiber membrane contactor

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(Membrane Contactor #1), as illustrated in Figure 1. The contactor was 6.7 cm in diameter by 28

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cm long and contained polypropylene fibers with a total external surface area of 1.4 m2. The

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membrane contactor was operated as a gas-liquid contactor and had its lumen side connected to a

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tank of CO2 gas. The effluent fitting on the lumen side was capped and the CO2 pressure inside

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the contactor was fixed at 104±4 kPa. Effluent solution from the gas-liquid contactor was then

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used to regenerate WBA media into HCO3- form. Effluent from the WBA media was combined

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with a base stream that bypassed the contactor. A series of 10 regeneration experiments were

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performed to assess how the fraction of the base stream that was fed through the membrane

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contactor affected the pH and carbonate content of the feed stream to the FBCR.

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Water Softening

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Water softening experiments were performed to determine ion balances between softening

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and regeneration cycles.

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synthetic tap water prepared in UPW with pH=7, containing: 6.4 meq/L Na+, 2.8 meq/L Ca2+,

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1.66 meq/L Mg2+, 4.73 meq/L Cl-, 3.46 meq/L SO42-, 2.67 meq/L HCO3-, 0.002 meq/L CO32- and

The ion exchange media was loaded with hardness ions using a


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0.53 mM H2CO3. This water composition contains the three most predominant cations and the

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four most predominant anions.17 Other cations typically found in water, such as K+, Cu2+, and

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Zn2+ are normally present at much lower concentrations, and are not considered here. The water

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composition was selected to be representative of a very hard water with a level of noncarbonate

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(permanent) hardness of 40%. This level of permanent hardness represents the high end of

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values that may be found in real waters, which typically have noncarbonate hardness

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concentrations that are 0 to 25% of the total hardness.18,19

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Because the WAC media was operated in H+ form, a column of WBA media in the HCO3-

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form was needed to provide additional bicarbonate to remove permanent hardness, as illustrated

211

in Figure S3 in the Supporting Information. The water softening system consisted of a column

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of Purolite A103 WBA media loaded with sulfate and bicarbonate followed by a column of

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Purolite C104 media operated in the H+ form. The WBA media had a manufacturer’s reported

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exchange capacity of 1.6 eq/L and the WAC media had a capacity of 4.0 eq/L. Because of its

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lower exchange capacity, the volume of WBA media (16.3 mL) was 2.5 times greater than the

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volume of WAC media (6.5 mL). The media were packed in glass columns with diameters of 1

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cm and were operated at flow rate of 0.7 mL/min, corresponding to an empty bed contact time of

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9.4 minutes for the WAC column. Two Gilson FC204 fraction collectors were used to collect

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effluent samples from each column every 60 minutes.

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Solution Analyses

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An Agilent model 5100 Synchronous vertical dual view inductively coupled plasma optical

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emission spectrophotometer (ICP-OES) was used to determine the concentration of sodium,

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calcium, magnesium, and sulfate in aqueous samples. Prior to analysis, samples were filtered

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with 0.45 µm syringe filters and acidified using HCl. Alkalinity was measured colorimetrically


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using the methyl orange method. 20 Solution pH values were measured using an AquaPro

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9156APWP epoxy body pH probe and a Symphony model SB90M5 multimeter.

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coefficients and saturation indices were calculated using the PHREEQC modeling package from

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the U.S. Geological Survey.21 The PHREEQC package calculates activity coefficients using the

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Davies or the extended WATEQ Debye-Huckel equation.22

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Results and Discussion

Activity

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The voltages required for acid and base generation were found to depend strongly on the

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pH of the feed solution. Figure 2 shows the voltage drop per unit cell for an 8 cell BMED stack

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using Tokuyama BP-1 bipolar membranes. Increasing pH values for the feed solution resulted in

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increasing voltage drops per unit cell for both NaCl and Na2SO4 solutions. This effect has not

235

been previously reported in the BMED literature, and cannot be attributed to higher ionic

236

conductivities of the lower pH solutions. As shown in a previous publication, >90% of the

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voltage drop for acid and base generation by this membrane in 100 mM salt solutions can be

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attributed to the water splitting reaction inside the bipolar membrane.23 Thus, the feed solution

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pH greatly affected the resistance for the water splitting reaction. Figure S4 in the Supporting

240

Information shows a similar effect for the Fumasep FBM membrane, which performed similarly

241

to the Tokuyama membrane for acid and base production. This effect of feed solution pH values

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on the voltage requirement for water splitting cannot be explained by current theories on BMED.

243

This issue notwithstanding, to minimize the voltage for operating the cell, all regeneration

244

experiments were performed with a feed solution pH value of ~2.

245

Two different flow configurations were used to generate acids and bases. In the first

246

configuration, the same 100 mM salt solution was fed into all chambers of the BMED stack.

247

This configuration resulted in SO42- concentrations in the acid chamber of ~200 mM. When


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sulfuric acid is used to regenerate ion exchange media, acid concentrations less than 200 mM are

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recommended in order to avoid CaSO4 precipitation in the ion exchange media.24 To minimize

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SO42- concentrations in the acid, a second configuration was investigated wherein the diluate

251

solution was fed into the acid and base chambers, as illustrated in Figure 1. This resulted in

252

SO42- levels in the acid that remained below 200 mM, even for the most concentrated acids that

253

were generated, as illustrated in Figure S5 in the Supporting information.

254

255 256 257 258

Figure 2. Voltage drop per unit cell (Δ Ψ ) as a function of current density (i) for an 8 cell stack of Tokuyama BP-1 membranes operated over a range in feed pH values using 100 mM Na2SO4 or NaCl solutions.

259 260

The acid or base concentration that could be generated was a function of both the

261

volumetric flow rate (Q), current density (i), area per BMED membrane (A), number of cells (n)

262

and current utilization ( ), as given by:

263

(1)

264

Figure 3 illustrates the current utilization for acid and base production, and the acid and base

265

concentrations, as a function of the current density. The acid concentrations were 20-25 mM


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higher than the base concentrations due to the ~10 meq/L acid concentration in the feed solutions.

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Current utilizations less than 1 resulted from acid and base leakage into the diluate compartments

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of the cell. Lower current utilizations for acid production at current densities below 20 mA/cm2

269

can be attributed to greater acid than base leakage.

270

271 272 273 274

Figure 3. Current utilization ( ) as a function of current density for 8 cell stack of Fumasep BPM operated with a pH=2 feed solution of 100 mM Na2SO4 to all chambers. Also shown is the acid and base concentration as function of current density.

275 276 277

Ion Exchange Regeneration The WAC ion exchange media was regenerated using acid with pH values ranging from 0.8

278

to 1 in either 100 mM Na2SO4 or NaCl electrolyte solutions.

279

experiments were performed using the same ion exchange cartridges, but varying the fraction of

280

the base stream that was fed through membrane contactor #1.

281

carbonate content of the feed stream to the FBCR, and also affected the amount of Ca2+ removal,

282

as shown in Table S1 in the Supporting Information.

283

concentrations from the WAC media are shown in Figure S6 in the supporting information. By


A total of 10 regeneration

This affected the pH and

Typical effluent hardness ion

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160 minutes elapsed, both Ca2+ and Mg2+ had been completely removed from WAC#1, while 99%

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of the Mg2+ and 53% of the Ca2+ had been removed from WAC#2.

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The regeneration stream was continuously fed into the FBCR along with a base stream

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containing dissolved carbon dioxide. Adding CO2 converts alkalinity due to hydroxide ions into

288

carbonate alkalinity.

289

calculated from the change in hydroxide ion concentration ([OH-]), according to:

290 291

292

293

Therefore, the final H2CO3, HCO3- and CO32- concentrations can be

[OH-]i - [OH-]f = [HCO3-]f + 2[CO32-]f

(2)

and the mass action expressions for the reactions: {

}{

}

{

(3)

}

{

}{ {

} }

(4)

294 295

The activity coefficients and final concentrations of the carbonate species exiting the membrane

296

contactor are shown in Table S2 for a range of operating conditions using the NaCl electrolyte.

297

Figure 4 shows the concentrations of hardness ions in the FCBR for the regeneration data

298

shown in Figure S6. In this experiment, 33% of the base stream was passed through membrane

299

contactor #1 and the pH of the carbonate solution fed into the reactor was 12.5. The pH remains

300

stable because the rate of OH- ion addition is equal to the equivalents of acidic protons in HCO3-

301

and H2CO3 that are also entering the reactor. Bicarbonate and carbonic acid will react with OH-

302

to form water and CO32- ions. Throughout the course of the regeneration process, hardness ion

303

concentrations in the FBCR remained below 0.07 mM. Based on the aqueous concentrations in

304

the FBCR, 99.7% of the Ca2+ and 95.8% of the Mg2+ removed from the WAC crystallized in the

305

reactor. Effluent from the FBCR was passed through a ceramic microfilter and sampled. The


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concentration of Ca2+ ions in the filtered effluent were 0.01 to 0.05 mM lower than in the reactor,

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indicating the presence of CaCO3 crystals not associated with the garnet sand.

308

Mg2+ concentrations in the filtered effluent ranged from 0.03 to 0.07 mM lower than in the

309

reactor. This also suggests the presence of suspended Mg(OH)2 crystals smaller than the ~0.1

310

µm filter cutoff. The near zero Mg2+ concentrations in the filtered effluent suggest that the filter

311

was more effective at filtering suspended Mg(OH)2 crystals than suspended CaCO3 crystals.

The

312

313 314 315 316 317 318

Figure 4. Hardness ion concentrations in the FBCR operated at a pH of 12.5 using 100 mM Na2SO4 as the electrolyte. Also shown are hardness ion concentrations after passing through a ceramic microfilter and after passing through WAC #3 ion exchange media. Influent hardness ion concentrations ranged from 0.1 to 30 mM and are shown in Figure S6 in the Supporting Information.

319 320 321 322

Saturation indices, defined as: (

)

(5)

323

where: IAP is the ion activity product, and Ksp is the mineral solubility product, were calculated

324

for the conditions as the end of the regeneration experiment, as shown in Table S3 in the

325

Supporting Information. Based on the measured ion concentrations in the reactor, the saturation

326

indices for CaCO3 (SI=3.62) and Mg(OH)2 (SI=74.1) were significantly greater than 1,


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indicating very high supersaturation of these species.

328

concentration measurements may have been impacted by crystals that were not removed by the

329

0.45 µm syringe filters. Thus, the hardness ion concentrations reported in Table S3 may be

330

greater than the dissolved hardness ion concentrations. Operating crystallization reactors at very

331

high saturation indices is undesirable and may promote homogeneous nucleation.10,11,12, 25 , 26

332

Furthermore, high levels of supersaturation are not needed to promote rapid crystallization.10,12,25

333

This indicates that the pH and total carbonate concentrations in the FBCR can likely be reduced

334

without adversely impacting its performance.

335

operating with NaCl electrolyte solutions, as shown in Figure S7 in the supporting information.

This suggests that hardness ion

Similar results were observed for systems

336

The filtered effluent from the FBCR was passed through WAC#3 media in the H+ form in

337

order to remove any residual hardness ions and to lower the pH of the effluent solution. This

338

process resulted in removal of less than 0.01 mM Ca2+ and less than 0.005 mM Mg2+. The small

339

amount of removal may be attributed to the low concentrations of hardness ions and the high

340

concentration of Na+ ions (200 mM), and to the presence of nanoparticulates passing through the

341

0.1 µm ceramic filter. Any nanoparticulates exiting WAC#3 would dissolve in the pH=2 feed

342

solution to the BMED stack. The low concentration of Mg2+ ions going back into the BMED

343

stack is not likely to cause a membrane fouling problem. For example, based on a base chamber

344

pH of 12.2, a Mg2+ concentration of 6.0 x 10-5 mM is required to reach saturation of Mg(OH)2 in

345

the base chamber. This is greater than the Mg2+ concentration of 4.5 x10-5 mM in WAC#3

346

effluent.

347

precipitation in the base chamber if carbonate concentrations in the feed stream are sufficiently

348

high. For example, based on a Ca2+ concentration in WAC#3 effluent of 6.8 x 10-3 mM, a CO32-

349

concentration greater than 0.67 mM in the base chamber would lead to supersaturation of CaCO3,

In contrast, the Ca2+ in the feed stream to the BMED stack may cause CaCO3


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assuming the Ca2+ concentration in the feed is the same as that in the base chamber. This is a

351

reasonable assumption given that cation electromigration into the base chamber will be

352

dominated by Na+ ions, since they are at a concentration that is ~8000 times greater than Ca2+

353

ions in WAC#3 effluent.

354

The CO32- and HCO3- concentrations being returned to the feed solution can be minimized

355

by acidifying the effluent from the FBCR and stripping the resultant CO2 into a high pH stream,

356

as illustrated in Figure 1. For the data in Figure 4, the pH of the solution entering WAC#3 was

357

12.5 and the carbonate alkalinity was 6.10 meq/L. The effluent from WAC#3 was at a pH value

358

of 5.4 and was fed into the tube side of membrane contactor #2, while a pH=12.2 solution was

359

fed into the lumen side of the contactor. This process reduced the carbonate alkalinity from 6.1

360

to 2.6 meq/L. Carbonate remaining in solution as H2CO3 was stripped out as CO2 gas after

361

entering the well-stirred feed tank with a pH=2. However, in order to prevent any CaCO3

362

precipitation in the base chamber with a feed Ca2+ concentration of 0.025 mM, H2CO3 levels in

363

the feed solution must remain below 0.67 mM.

364

Ion Balance

365

Operation of a closed-loop process using a fixed amount of regenerant solution requires

366

that ions are not lost from the system during the water softening step. Thus, the WAC media

367

must be operated in H+ form, rather than in Na+ form. WAC media in H+ form will only remove

368

temporary (i.e., carbonate) hardness, since HCO3- is necessary to react with the H+ released by

369

the media. In a typical water, the bicarbonate concentration in equivalent units represents from

370

75 to 100% of the total hardness.18 Thus, in most waters there is insufficient bicarbonate for

371

operating the WAC media in H+ form, and the alkalinity of the water will need to be augmented

372

by HCO3-.


17


Page 18 of 27

373

Figure 5 shows the performance of the softening system where a WBA column containing

374

HCO3- preceded the WAC softening column, as illustrated in Figure S3. Adding alkalinity to the

375

water via the WBA column was able to remove the permanent hardness. For the first 300 empty

376

bed volumes, more than 99.9% of the Mg2+ and more than 98.2% of the Ca2+ were removed.

377

This shows that the WBA regeneration process was able to put sufficient HCO3- on the WBA

378

media, despite the 5 times higher concentration of SO42- (100 mM), as compared to HCO3- (20

379

mM) in the regenerant solution. This result was not expected given that divalent ions are

380

normally preferred over monovalent ions on WBA media.27 The higher than expected uptake of

381

HCO3-can likely be attributed to a loss in ionic selectivity due to the high salt concentration in

382

the regenerant solutions.28 The effluent pH values from the WAC media were almost always

383

within ±1 pH unit of the neutral pH feed solution. Operating the WAC in H+ form with HCO3- in

384

the feed solution resulted in production of H2CO3 and slight acidification of the effluent solution.

385

Variation in the effluent pH values shown in Figure 5 can be attributed to loss of carbonic acid

386

via degassing CO2 from the samples before analysis.

387

Breakthrough experiments were also run without the WBA column preceding the WAC

388

media. These experiments showed immediate breakthrough of Ca2+ and Mg2+, and had effluent

389

pH values from the WAC media ranging from 2.5 to 5.7, as shown in Figure S8 in the

390

Supporting Information. Over the course of 800 empty bed volumes, the total hardness ion

391

removal was equivalent to 87% of the carbonate hardness. Thus, without the WAC loaded with

392

HCO3-, removal of hardness ions was less than the equivalents of HCO3- in the feed stream. This

393

can be attributed to removal of Na+ ions by the media that was equivalent to ~25% of the

394

hardness ions removed. Experiments were also performed with a column of WBA media in the

395

free base form preceding the WAC media. This was not effective, and the WAC media had


18

Page 19 of 27


396

immediate breakthrough of both Ca2+ and Mg2+, and effluent pH values ranging from 2.4 to 3.6,

397

as shown in Figure S9.

398 399

400 401 402

Figure 5. Effluent hardness ion concentrations and effluent pH values for ion exchange system consisting of WBA media in HCO3-/SO42- form, and WAC media operated in H+ form.

403 404

Removal of Na+ ions (here, a stand-in for all alkali metal ions) during the softening cycle

405

will lead to the accumulation of Na+ ions in the closed-loop regeneration system. Regenerating

406

the WAC using acid with a pH=1 completely removes all Na+ ions and releases them into the

407

regenerant solution. This increase in Na+ concentration will be accompanied by an equivalent

408

loss of H+ ions from the regeneration brine when the media is used again for water softening.

409

In addition to accumulating Na+ ions, SO42- and Cl- ions will also accumulate in the

410

regenerant solution. When the WBA media is regenerated as shown in Figure 1, 11% of the

411

media capacity is occupied by HCO3- and 89% by SO42-. When the WBA media is placed into

412

service, the HCO3- gets replaced by SO42- and Cl- from the feed stream. When the WBA media

413

is regenerated, these SO42- and Cl- ions are added to the regenerant solution. Balancing this


19


Page 20 of 27

414

increase in SO42- and Cl- concentrations is a loss of HCO3- from the regenerant solution when the

415

media is put back into service. The accumulation of Na+, SO42- and Cl- in the regenerant solution

416

will eventually require dilution of the regenerant solution in order to avoid CaSO4 precipitation

417

in the WAC media during regeneration. In addition, disposal of some of the regenerant solution

418

will be required. The simplest disposal option would be to add it to the softened water. This will

419

not increase the total dissolved solids load of the water as compared to the feed, since the Na+,

420

Cl- and SO42- ions were originally present in the feed water.

421

In addition to accumulating salt, the closed-loop system may also accumulate acid or base.

422

The BMED cell produces identical amounts of acid and base, and thus the performance of the

423

ion exchange media will determine acid or base accumulation in the regenerant solution.

424

Removal of one mole of Ca2+ ionsand precipitating it as CaCO3 (using CO2 gas as the source of

425

carbonate), consumes 2 equivalents of acid for ion exchange regeneration and consumes 2

426

equivalents of base for converting dissolved H2CO3 into CO32-. Similarly, removal one mole of

427

Mg2+ and precipitatitng it as Mg(OH)2, consumes 2 equivalents of acid and 2 equivalents of base.

428

However, to remove one mole of Na+ from the WAC and replace it with H+ consumes one

429

equivalent of acid and zero equivalents of base. Thus, more acid is consumed than base for ion

430

exchange sites occupied by Na+ ions during the softening step. However, for feed water with

431

noncarbonate hardness, base will be lost from the closed-loop system in order to provide

432

alkalinity for hardness removal. The main determining factor whether acid or base accumulates

433

is the difference between the fraction of the total hardness that is noncarbonate hardness, and the

434

fraction of the ion exchange sites on the WAC media that become occupied by Na+ ions during

435

the softening step.

436

equivalents of hardness removed can be expressed as:

For one regeneration cycle, the equivalents of acid accumulated per


20

Page 21 of 27


437

(6)

438 439

where V is the volume of water softened at hardness ion breakthrough, and the qi are the

440

equivalent concentrations for species i on the WAC media at breakthrough. In instances where

441

there is no noncarbonate hardness, the term:

442

equation (5), and the base accumulation will be equal to

, will be zero in .

443

In many practical circumstances, the accumulation of acid or base will be small. For

444

example, noncarbonate hardness values typically range from 0 to 25% of the total hardness.19

445

The relative loading of Na+ ions to hardness ions on WAC media can be determined from the

446

separation factor (

), defined as:28 [

]

[

]

447

(7)

448

and the aqueous equivalent concentrations of Na+, Ca2+ and Mg2+ ions in the water that is being

449

softened. The

450

recent national survey of major ion concentrations in drinking water reported equivalent

451

concentrations of Na+=1.65 meq/L, K+=0.13 meq/L, Ca2+=1.5 meq/L, and Mg2+=0.74 meq/L.17

452

Using these values, the ratio of alkali metal equivalents to hardness ion equivalents is 0.79. Thus,

453

a typical value for the fraction of the ion exchange media occupied by alkali metal ions can be

454

calculated as:

for the Purolite C104 media used here was determined to be 0.30. A

455

(8)

456

Using a noncarbonate hardness fraction of 0.12, that is in the middle of the typical range, and

457

alkali metal ion removal equal to 24% of the total hardness removal, equation 7 indicates base

458

accumulation in the system equivalent to 12% of the hardness removed. Slight accumulation of


21


Page 22 of 27

459

base may be beneficial, since excess base could be added to the softened water that was acidified

460

by the softening process. Alternatively, if acid accumulates and is added back into the softened

461

water, degassing of the water may be desirable to raise the pH.

462

Small amounts of water will also be lost from the system due to water splitting inside the

463

bipolar membrane. For water with 5 meq/L of hardness ions, ~5 meq of acid and base will be

464

needed per liter of water treated. In this case, producing acid and base via the splitting of water

465

represents a fractional water loss of 9.0 x10-5.

466

Energy Requirements

467

The practicality of the closed-loop regeneration process will depend on the energy required

468

for electrochemically producing the acid and base streams.

469

Information shows the required voltage per unit cell ( ) as a function of current density for

470

operating the BMED stack using 100 mM Na2SO4 as the feed solution. When the cell is

471

operated at a current density of 35 mA/cm2, the current utilization ( ) is the same for both acid

472

and base production, and is 75%. Thus, the energy required for producing 1 mole of acid or 1

473

mole of base can be calculated from the ratio of the applied power (

474

rate of acid or base (

Figure S10 in the Supporting

) to the molar production

), as given by:

475

(9)

476

where I is the applied current, n is the number of unit cells in the BMED stack, and F is

477

Faraday’s constant. Since they are produced simultaneously, equation 8 ascribes half the energy

478

for acid production and half for base production. For operating the cell at a current density of 35

479

mA/cm2, the applied voltage per unit cell is 4 V, and the energy required per mole of acid or base

480

is 0.071 kWh. Assuming an electrical energy cost of $0.07/kWh29, the cost for 1 mole of acid


22

Page 23 of 27


481

plus 1 mole of base is $0.010. This is more than 15 times lower than the total cost of $0.155 for

482

purchasing acid and base from a local supplier in 180 lb containers (Hill Brothers Chemical Co.,

483

Tucson), where the cost per mole of 35% HCl is $0.069 and 99% NaOH is $0.086. In addition

484

to the energy cost to operate the cell, energy will be required for pumping the solutions through

485

the cell. Past experience shows that the energy required for this is similar to the energy required

486

to operate the BMED stack.5 Thus, even when these costs are included, the energy costs for

487

producing the acids and bases are approximately an order of magnitude smaller than those for

488

purchasing acids and bases.

489

Scalability

490

The membrane area needed per volume of water softened can be estimated from data in this

491

study. The stoichiometric amount of acid and base required for treating 1000 m3/d of water with

492

5 meq/L of hardness is 5000 eq/d for removing hardness minerals from the WAC. In addition,

493

the WAC will remove alkali metal ions equal to ~25% of the hardness ions removed. Thus, the

494

amount of acid and base required will be ~6250 eq/d. A BMED cell with 1 m2 of membrane area

495

operating at 35 mA/cm2 with

496

m2 of BMED membrane area is required to produce 6250 eq of acid and base per day. This can

497

be accommodated in one stack with 107 unit cells with membranes that are 50 cm by 50 cm in

498

dimension. This is within the range of commercial stacks that may have up to 300 unit cells.5

= 0.75 produces 235 eq of acid and base per day. Therefore, 26.7

499

Acid and base wastage in the system will increase the required membrane area over the

500

stoichiometric amount. In order to operate the system as a closed loop, acid is required to reduce

501

the pH of the solutions exiting the FBCR. If the FBCR is operated at a pH of 11, and the

502

concentration of the carbonate species are each 1 mM, 14 meq/L of acid is required to remove

503

CO32- and HCO3- (as CO2), and to acidify the solution to a pH value of 2. For a system operating


23


Page 24 of 27

504

at 35 mA/cm2, the acid generation rate is 225 meq/L when operated at the flow rates used here

505

(Figure 3). Thus, continuous acid and base consumption of 14 meq/L associated with removing

506

carbonate and acidifying the FBCR effluent would only increase the required acid and

507

production by 6.2% over the stoichiometric amount needed for ion exchange regeneration.

508 509

Acknowledgements

510

The National Science Foundation (CBET-1235596) provided funding for this work

511 512

Conflict of Interest Disclosure

513

The authors declare no competing financial interest.

514 515

Supporting Information

516

Figures illustrating: the experimental apparatus, the configuration of the water softening system,

517

a bipolar membrane unit cell, the voltage drop per unit cell, sulfate concentrations in acid stream

518

as a function of current density, hardness ion concentrations in WAC effluent and in FBCR

519

versus elapsed regeneration time, effluent hardness ion concentrations and pH for WAC media

520

operated in H+ form, effluent hardness ion concentrations and pH for WAC media operated in H+

521

form with preceding WBA media operated in free base form, voltage drop per unit cell versus

522

current density, and Tables illustrating membrane contactor performance and levels of hardness

523

ion removal and hardness ion concentrations in the FBCR are provided in the Supporting

524

Information. This material is available free of charge via the Internet at http://pubs.acs.org.

525


24

Page 25 of 27

TOC Art

X-

-

OH

M+

Acid + + + + + + + + + + + + + + + + +

-

+

H

DI + + + + + + + + + + + + + + + + +

+

M

X-

Base -

-

OH

Acid + + + + + + + + + + + + + + + + +

-

+

H

X-

+ + + + + + + + + + + + + + + + +

M+

Cathode

Anode

M

-

Water Splitting Catalyst Layer

Base +

Water Splitting Catalyst Layer

526


X-

Salt Solution

527 528


25


529

Page 26 of 27

References

(1) Water Quality and Treatment: a Handbook of Community Water supplies, 5th ed.; McGraw Hill: New York, 1999. (2) Management of water treatment plant residuals; American Water Works Association: Denver, CO, 1996. (3) Central Arizona Salinity Study, Phase II, Final Report; U.S. Bureau of Reclamation, 2006. (4) Santa Clara River Chloride Reduction Ordinance of 2008; Santa Clarita Valley Sanitation District of Los Angeles County, 2008. (5) Strathmann, H. Ion-Exchange Membrane Separation Processes; Elsevier: Amsterdam, 2004. (6) Badruzzaman, M.; Oppenheimer, J.; Adham, S.; Kumar, M. Innovative beneficial reuse of reverse osmosis concentrate using bipolar membrane electrodialysis and electrochlorination processes. J. Memb. Sci. 2009, 326, 392–399. (7) Chiao, Y. C.; Chlanda, F. P.; Mani, K. N. Bipolar membranes for purification of acids and bases. J. Memb. Sci. 1991, 61, 239–252. (8) Mani, K. N.; Chlanda, F. P.; Byszewski, C. H. Aquatech membrane technology for recovery of acid/base values for salt streams. Desalination 1988, 68, 149–166. (9) Onsager, L. Deviations from Ohm’s law in weak electrolytes. J. Chem. Phys. 1934, 2, 599615. (10) Graveland, A.; van Dijk, J. C.; de Moel, P. J.; Oomen, J. H. C. M. Developments in water softening by means of pellet reactors. J. Am. Water Works Assoc. 1983, 75 (12), 619-625. (11) Harms, W. D.; Robinson, R. B. Softening by fluidized bed crystallizers. J. Env. Eng. 1992, 118, 513–529. (12) Van der Veen, C.; Graveland, A. Central softening by crystallization in a fluidized bed process. J. Am. Water Works Assoc. 1988, 80 (6), 51-58. (13) Stanford, B.D.; Leising, J.F.; Bond, R.G.; Snyder, S.A. Inland desalination: Current practices, environmental implications, and case studies in Las Vegas, NV. Sustainability Sci. Eng. 2010, 2, 327-350. (14) Mahvi, A. H.; Shafiee, F.; Naddafi, K. Feasibility study of crystallization process for water softening in a pellet reactor. Int. J. Environ. Sci. Technol. 2005, 1 (4), 301-304. (15) van Schagen, K.M.; Rietveld, L.C.; Babuska, R.; Kramer, O.J. Model-based operational constraints for fluidised bed crystallisation. Water Res. 2008, 42, 327–337. (16) Giesen, A.; Erwee, H.; Wilson, R.; Botha, M.; Fourie, S. Experience with crystallization as sustainable, zero-waste technology for treatment of wastewater. In Proceedings of the International Mine Water Conference; Pretoria, S Africa, Oct 19-23, 2009; ISBN 978-009802623-5-3.


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(17) Patterson, K.Y.; Pehrsson, P.R.; Perry, C.R. The mineral content of tap water in United States Households. J. Food Compost. Anal. 2013, 31 (1), 46-50. (18) Davis, M. Water and Wastewater Engineering; McGraw-Hill: New York, 2010. (19) Mechenich, C.; Andrews, E. Home Water Safety: Interpreting Drinking Water Test Results; University of Wisconsin Cooperative Extension Publication G3558-4, 2004 (20) Methods for the Chemical Analysis of Water and Wastes (MCAWW) (EPA/600/4-79/020) (21) Parkhurst, D.L.; Appelo, C.A.J. Description of input and examples for PHREEQC version 3: a computer program for speciation, batch-reaction, one-dimensional transport, and inverse geochemical calculations. U.S. Geol. Surv. 2013, No. 6-A43. (22) Parkhurst, D.L.; Appelo. C.A.J. User's guide to PHREEQC (Version 2): A computer program for speciation, batch-reaction, one-dimensional transport, and inverse geochemical calculations. U.S. Geol. Surv. 1999. (23) Davis, J.R.; Chen, Y.; Baygents, J.C.; Farrell, J. Production of acids and bases for ion exchange regeneration from dilute salt solutions using bipolar membrane electrodialysis. ACS Sustainable Chem. Eng. 2015, 3 (9), 2337-2342. (24) Bornak, W. Ion Exchange Deionization for Industrial Users; Tall Oaks Publishing: Littleton, CO, 2003. (25) Tai, C.Y.; Chien, W.C.; Chen, C.Y. Crystal growth kinetics of calcite in a dense fluidizedbed crystallizer. A.I.Ch.E.J. 1999, 45, 1605–1614. (26) van Schagen, K.M.; Rietveld, L.C.; Babuska, R.; Kramer, O.J. Model-based operational constraints for fluidised bed crystallisation. Water Res. 2008, 42, 327–337. (27) Clifford, D.; Weber, W. J. The determinants of divalent/monovalent selectivity in anion exchangers. React. Polym. 1983, 1, 77-89. (28) Clifford, D. Ion Exchange and Inorganic Adsorption in Water Quality and Treatment, 5th ed.; R. D. Letterman ed.; American Water Works Association, McGraw-Hill: New York, 1999; pp 9.1-9.91. (29) Eia. gov., Electricity Data Browser; http://www.eia.gov/electricity/data/browser/#/topic/7?agg=2,0,1&geo=g&freq=M (accessed Dec. 5, 2014).


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Electrochemical Ion-Exchange Regeneration and Fluidized Bed

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