Diffusion dialysis-concept, principle and applications

Journal of Membrane Science 366 (2011) 1–16

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Journal of Membrane Science journal homepage: www.elsevier.com/locate/memsci

Diffusion dialysis-concept, principle and applications Jingyi Luo a , Cuiming Wu b , Tongwen Xu a,∗ , Yonghui Wu a a b

Lab of Functional Membranes, School of Chemistry and Material Science, University of Science and Technology of China, Hefei 230026, PR China School of Chemical Engineering, Hefei University of Technology, Hefei 230009, PR China

a r t i c l e

i n f o

Article history: Received 25 July 2010 Received in revised form 9 October 2010 Accepted 12 October 2010 Available online 16 October 2010 Keywords: Diffusion dialysis Ion exchange membrane Acid recovery Alkali recovery Integrated membrane processes

a b s t r a c t Diffusion dialysis (DD) is an ion-exchange membrane (IEM) separation process driven by concentration gradient and has been applied for separation and recovery of acid/alkali waste solutions in a cost-effective and environmentally friendly manner. This review of DD covers the principles, models, applications (strong acid/weak acid/alkali separation and recovery), and its integration with other techniques, such as electrodialysis, ion exchange membrane-electrowinning, continuous membrane extraction, vacuum membrane distillation, and ceramic membrane micro-filtration. Notably, different factors including properties of the membranes, nature of the waste solution and running conditions are discussed and correlated with the DD performances. © 2010 Elsevier B.V. All rights reserved.

Contents 1. 2.

3. 4.

5.

6.

7.

Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . General description of diffusion dialysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 2.1. Principles of diffusion dialysis. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 2.2. Models for diffusion dialysis process . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 2.3. Experimental setups . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Membranes for diffusion dialysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Inorganic acid recovery using diffusion dialysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.1. Sulfuric acid (H2 SO4 ) recovery . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.1.1. The effect of anion exchange membrane (AEM) . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.1.2. The effect of physico-chemical properties of the waste acidic solution . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.1.3. The effect of acid/salt concentration, flow rate/ratio, temperature and other factors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.2. Hydrochloric acid (HCl) recovery . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.2.1. The salt effect in hydrochloric acid (HCl) recovery . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.2.2. The effect of complexes on hydrochloric acid (HCl) recovery . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.2.3. The effect of other factors on hydrochloric acid (HCl) recovery . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.3. Nitric acid (HNO3 ) recovery . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4.4. Schematic and prototypical designs for diffusion dialysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Treatment of organic acid using diffusion dialysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 5.1. The factors determining the separation of weak acid from its salt . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 5.2. Transport characteristics (diffusivity, permeability, mass transfer coefficient and overall dialysis coefficients) . . . . . . . . . . . . . . . . . . . . . . . . . . Base recovery using diffusion dialysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 6.1. Alkali waste and treatment . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 6.2. The application and development of diffusion dialysis for alkali separation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Integrated processes based on diffusion dialysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7.1. Diffusion dialysis integrated with electrodialysis . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7.2. Diffusion dialysis integrated with IEM-electrowinning . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

∗ Corresponding author. Tel.: +86 551 360 1587; fax: +86 0551 3602171. E-mail address: [email protected] (T. Xu). 0376-7388/$ – see front matter © 2010 Elsevier B.V. All rights reserved. doi:10.1016/j.memsci.2010.10.028

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7.3. Diffusion dialysis integrated with continuous membrane extraction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7.4. Diffusion dialysis integrated with vacuum membrane distillation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7.5. Diffusion dialysis integrated with ceramic membrane microfiltration . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Summary and perspective . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Acknowledgements . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . References . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

1. Introduction Diffusion dialysis (DD) is an ion-exchange membrane (IEM) separation process driven by concentration gradient and is also known as concentration or natural dialysis [1]. The concept of dialysis was firstly proposed by Graham in 1861, as a way of separating relatively small molecules from large ones with a semi-permeable membrane [2]. Since the driving force for the separation process is mainly concentration gradient, dialysis is known as a spontaneous separation process [2]. Specifically, if IEMs (also a kind of semi-permeable membrane) are employed for dialysis, the process is defined as diffusion dialysis (DD). In the 1950s, the first diffusion dialyzer was invented [1], and more than 30 years later DD technique was first developed into an industrial membrane process in Japan [3,4]. As a spontaneous process, the process of DD gives rise to an increase in entropy and decrease in Gibbs free energy, so it is thermodynamically favorable. In comparison with some conventional processes, DD demonstrates a significant superiority [5–7]:

13 14 14 14 15 15

Nevertheless, DD also has its limitations, such as relatively low processing capability and efficiency. For instance, a diffusion dialyzer with an effective membrane area of 500 m2 can only dispose 58 m3 of waste acid for 1 day. And the concentration of recovered solution is limited by the equilibrium [1]. Consequently, DD is considered less important than some other membrane separation processes, such as electrodialysis and reverse osmosis. However, the unique advantages of DD, especially the environmental benignity and low energy consumption, may make DD more competitive as environmental pollution and energy shortage become more serious. Furthermore, with the improvement on IEMs and diffusion dialyzers, the processing capability and separation efficiency of DD can be substantially enhanced, and DD can be developed into a practical separation technique in the near future. This paper will give an overview of DD, including the general principles, the membranes used for DD, the applications of DD in inorganic acid/organic acid/alkali recovery, and the integrations of DD with other membrane processes. 2. General description of diffusion dialysis

(1) Higher efficiency in purifying wastewater; improvement on the productivity and quality of products. (2) Low energy consumption (DD runs under normal pressure and has no state change during the process, so no power is needed for running DD). (3) Low installation and operating cost; stable, reliable, and easy for operation. (4) No pollution of the environment. To date, DD has been successfully applied for recovery of acids and alkalis from the discharges from steel production, metal-refining, electroplating, cation exchange resin regeneration, non-ferrous metal smelting, aluminum etching, and tungsten ore smelting [2,3,8–11]. When it comes to associated studies, there has been an increased interest from institution-based research (Fig. 1).

Fig. 1. Chronology of diffusion dialysis documents. Source: www.scopus.com [search settings: TITLE-ABS-KEY (diffusion dialysis). Search date: July 9, 2010].

2.1. Principles of diffusion dialysis During the DD process, the ion transport is driven mainly by the concentration gradient, with observation of the Donnan criteria of co-ion rejection and preservation of electrical neutrality [2]. The separations of HCl and NaOH from their feed solution are illustrated in Figs. 2 and 3 to describe the principle of diffusion dialysis (DD). As shown in Fig. 2, HCl and its metal salts in the feed solution tend to transport to the water side due to the concentration difference across the membrane. Because of the presence of the AEM, the Cl− ions (or SO4 2− , NO3 − , PO4 3− , etc.) are permitted passage, while the metals in the waste solution are much less likely to pass. The H+ ions, although positively charged, have higher competition in diffusion than metal ions because of their smaller size, lower valence state and higher mobility. Hence they can diffuse along with the

Fig. 2. Illustration of the diffusion dialysis principle through the HCl separation process from its feed solution.

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force in this zone [16]. From combination of the three-phase membrane model with the principle of DD, we may roughly deduce that the DD process for acid–salt separation is realized as following: under the influence of concentration gradient, acidic anions have the priority to pass through the active region of membranes via the hopping mechanism, leading to the requirement of electrical neutrality in the water side. In order to satisfy with the electro neutrality requirement, cations (hydrated hydrogen) leak through the interstitial zone of membranes. The transport processes above make the separation process proceed continuously. As for the system with CEM for alkali recovery, the principle is analogous to that of acid. 2.3. Experimental setups

Fig. 3. Illustration of the diffusion dialysis principle through the NaOH separation process from Na2 WO4 solution.

Cl− ions (or SO4 2− , NO3 − , PO4 3− , etc.) to meet the requirement of electrical neutrality. The H+ transport is a key to the DD process [2,5,12]. Proper properties of the AEM are also necessary, including stability in acidic solution, high H+ permeability, strong rejection for other metal ions, relatively high water uptake (WR ) and poor water permeability. Separation process of NaOH from its feed solution (Na2 WO4 as an example) is illustrated in Fig. 3. NaOH and Na2 WO4 tend to transport to the water side due to the concentration difference across the membrane. Because of the presence of a CEM, the Na+ in the feed are permitted passage, while the WO4 2− ions are much less likely to pass through the membrane. Similar to H+ through an AEM, the hydroxyl ions (OH− ) have higher competition in diffusion than WO4 2− ions and can diffuse along with Na+ ions to meet the requirement of electrical neutrality. The OH− transport is also a key to the process [13], and the CEM with high stability in alkali solution, high OH− permeability, strong rejection for other anions, relatively high WR but poor water permeability are also required.

DD process can be realized through two kinds of runs: one is batch dialysis and the other is continuous dialysis as shown in Fig. 4(a) and (b): In the batch dialysis (Fig. 4(a)), a two-compartment cell of equal volume is separated by an IEM. The two compartments are filled with feed solution and water, respectively, and stirred at identical rates to minimize concentration polarization. The transport characteristics are determined from time dependences of the component concentration in both the compartments [17]. The batch dialysis is mostly used for experimental research in the lab, such as the determination of the acid diffusivity and separation factor, and it can be considered as a basic of the continuous dialysis. In the continuous dialysis, take the acid recovery process for an example (Fig. 4(b)), the diffusion dialyzer is separated by a number of IEMs into dialysate cells and diffusate cells, through which the feed and stripping water pass respectively in counter current direction. The transport characteristics are calculated from the concentrations and volumetric (or mass) liquid flow rates of the streams entering and leaving the dialyzer at steady state [18]. The data obtained in a continuous dialysis can be more valuable for practical reference, because the experimental conditions are more similar to those in practical dialysis unit. 3. Membranes for diffusion dialysis

2.2. Models for diffusion dialysis process Until now, two models have been used to describe DD permeation process. The first is the solution-diffusion model [14], which was put forward 20 years ago and has been accepted mainly for explanation of the transport in dialysis, reverse osmosis, gas permeation, and pervaporation [14]. The model can also be applied in DD separation process of strong acids and alkalis. According to this model, components dissolve in the membrane phase and then diffuse through the membrane down a concentration gradient. The separation is achieved due to their difference in dissolubility and diffusion rate [14]. This model for organic acids DD process may be less accurate since the dissociation degree of organic acids is low and their DD transport mechanism inside the membrane can be very complex [15]. The other and more popular model is three-phase membrane model [16], in which the membrane is assumably divided into three phases, i.e. a hydrophobic polymer, an active region (including the exchange fixed sites and the counter-ions) and an interstitial region. The water, which is indispensable for the migration of ions, exists mainly in the active and interstitial zones. The ions can transport through these two regions via different mechanisms: the anions in the active zone may move by a hopping mechanism, transferring from one exchange site to another; while hydrated protons and other cations prefer a “dragging” mechanism (restricted diffusion) in the interstitial zone, because there is no serious repulsive

As mentioned above, the membranes for DD can be either an AEM or a CEM. The former is for the separation of acids with the corresponding salts and the latter is for the similar separation of base mixtures. Due to the higher demands of acid recovery, more attention has been placed on AEMs compared to CEMs, and as such an increasing number of AEMs have been developed for DD applications recently. For instance, a series of plat AEMs and hollow fiber AEMs have been prepared from poly (2,6-dimethyl-1,4-phenylene oxide) (PPO) in our laboratory with quaternary amine [5]. Their properties (ion-exchange capacity (IEC), WR , etc.) can be adjusted by bromine substitution content and position, functional process as well as the amine- or silane-crosslinking degree. The effects of the membranes’ physico-chemical properties on their DD performances have also been investigated, and excellent DD effects can be observed [8,19–22]. CEMs for DD are relatively seldom reported compared with AEMs. In the work of Kiyono et al. [23], heterogeneous hollow fiber membranes with sulfonic acid groups are prepared using a wet spinning technique. The DD tests in the systems of NaOH + NaCl and NaOH + Na2 SO4 have been investigated, and relatively desirable separation effects have been achieved for some of the membranes [23]. Moreover, other plat CEMs such as hydrophilic fluorine membrane, hydrophobic fluorine membrane with perfluorinated sulfonic acid groups [24], and polyethylene heterogeneous mem-

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Fig. 4. Experimental set-up for diffusion dialysis: (a) batch dialysis and (b) continuous dialysis for the acid recovery.

brane with sodium sulfonate groups [25] have also been studied as matrices in recycling alkali from discharges [26]. Besides the membranes mentioned above, there are also commercially available membranes for DD process. Some typical examples are summarized in Table 1 [5,27–29] (www.ameridia.com/html/mbt.html; www.ensh.Dz; www.twmembrane.com).

4. Inorganic acid recovery using diffusion dialysis Sulphuric acid (H2 SO4 ), hydrochloric acid (HCl) or a combination of hydrofluoric and nitric acids (HF + HNO3 ) is often used as pickling agents in industries for a variety of metals etching and stripping processes, such as steel production, metal refining [30], non-ferrous metal smelting [31]. Large quantities of spent liquor are produced during the pickling steps. For example, 2 × 105 to 4 × 105 kg or ∼120 m3 of H2 SO4 waste is produced when 1 × 103 kg titanium white or vanadium is manufactured respectively [19]. And 2 × 104 to 4 × 104 kg HF + HNO3 waste is produced when one ton of titanium material is processed [20]. Though the acid liquor can be circulated, the accumulation of metal ions (especially metal ions with high valence) in the solution will result in a decreased efficiency of the pickling agent. And thus the eventual treatment is necessary [20,31].

Until now, different methods dealing with the acidic discharges have been applied, including [18,32]: cooling and crystallization, thermal decomposition, evaporation and crystallization, ion exchange, solvent extraction, distillation and electric-membrane separation methods, as well as direct disposal and neutralization with alkalis. Nevertheless, some native shortcomings still hinder their further developments and applications [18,32], such as high investment for equipment, large consumption of energy and alkalis, and environmental pollutions. DD process is developed accordingly, and the DD researches and applications for the recovery of specific acid will be elucidated in next section.

4.1. Sulfuric acid (H2 SO4 ) recovery Sulfuric acid (H2 SO4 ) is an important and relatively low-cost acid. When utilized for electroplating process and surface treatment of steel, waste solution containing free H2 SO4 and metallic ions (Fe(II), Ni(II), Cu(II), Zn(II), etc.) is generated [7,31]. Adoption of DD technique in those wastes treatment can not only recover H2 SO4 , but also reject the salts, and hence provide much industrial and environmental benefit. The recovery and rejection rates during DD process are determined by different factors, such as the membrane properties, the feed compositions, and the operation parameters, as we will further elaborate.

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Table 1 Properties of some commercially available ion exchange membranes for diffusion dialysis processes. Membrane

Type

Thickness [mm]

IEC [meq/g]

Area resistance [ cm2 ]

Materials

Manufacturer

Selemion DSV Neosepta AFN

AEM AEM

0.12–2.5 0.15–0.18

4.5–5.5 2.0–3.5

– 0.2–1.0

Asahi Glass, Tokyo, Japan Tokuyama Co., Japan

Neosepta AFX

AEM

0.14–0.17

1.5–2.0

0.7–1.5

Neosepta AMH

AEM

2.25–0.26

1.3–1.5

11.0–13.0

SB-6407 DF120-I

AEM AEM

0.152 0.23–0.32

2.15 1.9–2.2

0.3–1.2 1.5–2.1

Aminated polysulfone Polystyrene crosslinking and amination Poly-styrene-co-divinylbenzene, aminated Poly-styrene-co-divinylbenzene, aminated Aminated polysulfone BPPO amination

DF120-III

AEM

0.20–0.23

1.7–1.9

3.5–4.0

BPPO amination and crosslinking

DC120

CEM

0.20–0.25

1.5–2.0

1.5–2.0

Sulphonated PPO

4.1.1. The effect of anion exchange membrane (AEM) AEMs are the core of DD process and variation of their properties, including structure, charged groups and thickness, always affect the DD performances significantly. For instance, Ersoz et al. [33] compared the possible use of SB-6407 and Neosepta-AMH membranes in recovering acids by DD. SB-6407 membrane is a strongly basic AEM filter with quarternary ammonium groups, while AMH is a strongly basic AEM with −NC7 H7 functional groups. And a higher degree of acid separation effect is achieved by SB-6407 [33]. Though the authors provided no explanation for such difference in DD performances, some clues can be found from the differences of their IEC and thickness-SB-6407: 2.15 mmol/g, 0.152 mm; AMN membrane: 1.90 mmol/g, 0.26 mm [34]. We have also previously investigated the influences of PPObased membrane structure on membrane properties and their DD performances [19]. In our study, the benzyl and aryl groups of poly (2,6-dimethyl-1,4-phenylene oxide) (PPO), as shown in Fig. 5 can be brominated to different degrees, and surface-cross-linking degree can be varied by controlling the amination time. With benzyl bromine content (BBC) increasing, IEC and WR are enhanced and CR (fixed group concentration) is decreased. Meanwhile, with the increasing of aryl bromine content (ABC), WR is gradually decreased and IEC remains approximately unchanged, resulting in the increased CR . And with increasing in cross-linking degree, the membranes become more compact and hydrophobic, causing decreased WR and increased CR . DD tests of the membranes show that higher BBC improves the acid recovery, while higher ABC and cross-linking degree enhance the selectivity [19]. Correlation of the membrane properties and their DD performances with the DD principle and three-phase model draws the following conclusion: WR and CR of the membranes are two key parameters affecting DD effect – the higher WR favors the diffusion of ions, while the lower CR decreases electrostatic repulsion of the membrane to cations. CH3 O

CH3 Br2

O

Br

CH3

N(CH3)3

Aryl bromine Benzyl bromine (BPPO)

(PPO) CH3 O

Br

n CH2Br

CH3

CH3

+

n CH2N(CH3)3Br

O

O Br

CH2

N

H2C

Surface crosslinking of BPPO Fig. 5. The bromination and crosslinking reactions of PPO.

Br

Tokuyama Co., Japan Tokuyama Co., Japan Gelman Sciences Shandong Tianwei Membrane Technology Co., China Shandong Tianwei Membrane Technology Co., China Shandong Tianwei Membrane Technology Co., China

Besides the selectivity and permeability of ions, water osmosis is also influenced by the membrane properties, especially the membrane hydrophilicity [8]. PPO-based membrane can also be taken as the illustration: with high cross-linking degree and low WR value, the water osmosis is depressed significantly. The recycling burden and pumping energy, as well as the decrease of feed concentration is suppressed. This is especially meaningful for the DD system with low feed concentration, because the driving force for diffusion is intrinsically low for such system and further decrease of the feed concentration by water osmosis will render the diffusion process difficult. With PPO-based membrane from proper cross-linking time (8 h), the DD result of H2 SO4 /NiSO4 system is as following: water osmosis is relatively low (volume expansion factor of feed = 1.1), acid recovery ratio high (77%), nickel leakage ratio low (4%) and the selectivity relatively high (75%) [8].

4.1.2. The effect of physico-chemical properties of the waste acidic solution Different ionic forms and complexes may be present in the waste acidic solution, and the interaction between themselves (or between them and membranes) is also varied. Both can influence the transport rates of components and correspondingly affect the DD performances. In the system of H2 SO4 + CuSO4 [35], the competitive absorption of H2 SO4 and CuSO4 and the change of complex concentration have been utilized for explanation of the DD results: with increasing H2 SO4 concentration in feed, adsorbed CuSO4 concentration in the membrane, and the concentration of non-dissociated form of CuSO4 ((CuSO4 )non ) decrease, while Cu2+ ions concentration increase. Permeability of Cu2+ is lower than that of (CuSO4 )non . Hence the flux of CuSO4 decreases, and the rejection coefficient increases [35]. For the system of H2 SO4 + ZnSO4 [36], the sorption isotherm, the flux and the rejection rate of ZnSO4 follow the same trend as the system of H2 SO4 + CuSO4 . The difference is that more complexes bearing negative charge are supposed to be present in the membrane: [Zn(SO4 )2 2− , Zn(SO4 )3 4− and Zn(SO4 )4 6− ]. Those complexes have higher permeability and their concentrations decrease with the increase of acid concentration. Therefore, the flux of ZnSO4 decreases and the rejection coefficient increases [36]. Meanwhile, from the work of Azzeddine et al. on H2 SO4 + ZnSO4 and H2 SO4 + Na2 SO4 systems [37], ion pair formation of sulfate salt is assumed as the main reason for the decreased sulfate mobility in the membrane. The DD experiment, membrane resistance (Rm ) and sorption experiment results are in accordance with the observation of low sulfate mobility in the membrane: the permeability of sulfate is low and does not exhibit any significant effect on acid flux; the Rm treated with Na2 SO4 and ZnSO4 decrease less significantly

6

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than that with H2 SO4 ; the adsorbed amount of Na+ and Zn2+ is less than that of H+ [37]. Besides the effect of the complexes, the characteristics of the ions can also affect their permeability. For example, for the system of H2 SO4 + FeSO4 + VOSO4 , VO2+ cation rejection is always slightly higher than Fe2+ ion rejection due to the larger size of VO2+ [32]. 4.1.3. The effect of acid/salt concentration, flow rate/ratio, temperature and other factors Sulfuric acid concentration, salt sulfate concentration and operation temperature also affect DD performance significantly. For the continuous dialysis, the effect of the flow rate, flow ratio and number of AEMs should also be considered. The systems of H2 SO4 + CuSO4 , H2 SO4 + ZnSO4 , H2 SO4 + FeSO4 can be taken for illustrations. With the H2 SO4 concentration increasing, the absorption and permeability of sulphates (such as CuSO4 and ZnSO4 ) generally decreases [35,36]. The influence of increasing sulphates concentrations on their permeability is different under different conditions. For instance, in the H2 SO4 + FeSO4 system [32], increasing the FeSO4 concentration causes a decrease in its mobility due to the increased ionic strength and an increase in the concentration gradient. Therefore, the net increase in FeSO4 permeability is determined by the mutual interaction of ionic strength and concentration gradient [32]. The effect of sulphate concentration on acid recovery, namely the “salt effect” is not significant and nearly cannot be observed for most of the H2 SO4 + sulphates systems. This is somewhat odd and different from the observations of HCl systems (as will be introduced in the following Section 4.2.1). For instance, in the DD systems of H2 SO4 with FeSO4 , NiSO4 , CuSO4 , or ZnSO4 , enhancement of the H+ permeability by increase of sulphates can only be clearly observed in H2 SO4 –FeSO4 system at low FeSO4 concentrations (<0.5 mol/L) [7,19,37]. The influence of ZnSO4 or CuSO4 on H2 SO4 absorption or permeability is insignificant [38], and NiSO4 nearly has no effect on H2 SO4 recovery [8]. Absence of the “salt effect” can also be observed in other systems with metal ions of +1, +3 or +4 valences. Also, in the H2 SO4 waste containing Na2 SO4 , the presence of Na+ nearly does not change the flux of H2 SO4 in the whole investigated concentration range (0–2 mol/L) [37]; moreover, if Al2 (SO4 )3 exists in the waste solution, the H+ mobility is apparently retarded [38]. In the case of titanium ions (based on TiO2 , 0.25 mol/L H+ ), the enhancement of H2 SO4 permeability nearly cannot be observed in the investigated concentration range of Ti (0.1–1.4 mol/L) [19]; similar phenomenon is also observed in the case of VOSO4 with concentration range 0.03–0.09 mol/L [32]. The above observations may be interrelated with the ionic strength of the solution, complexes formed from the SO4 2− and metal ions, and the ion pair formations within the membrane. For the choice of running temperature for DD process, different factors should be taken into account. On one hand, the waste acid solution from a production line usually has a high temperature (much higher than 40 ◦ C). Higher temperature is also advantageous for acid recovery due to the increase of diffusion coefficient [21,39]. For example, increasing the temperature from 10 ◦ C to 30 ◦ C can increase the recovery of H2 SO4 from 44% to 63% [7]. On the other hand, the thermo-chemical stability of the AEM is limited and hence an operating temperature above 40 ◦ C is to be avoided. Therefore, a cooling system is necessary to decrease the feed solution to 25 ◦ C or lower in order to increase the life of the membrane [40]. As for the continuous dialysis, although the H2 SO4 recovery increases with the flow ratio of water to feed, the concentration of recovered H2 SO4 decreases [7,41]. When the flow ratio is constant, the recovered H2 SO4 concentration and its recovery increases firstly and then decreases with increasing flow rate [7,41]. Thus, suitable flow ratio and flow rate should be chosen. For instance, in the system of H2 SO4 + Al2 (SO4 )3 , about 80% H2 SO4 can be recov-

ered from the waste solution (concentration of H2 SO4 = 4.5 mol/L) at the flow rate of 0.26 × 10−3 m3 h−1 m−2 and flow ratio of 1:1. Under these conditions the concentration of recovered sulfuric acid is 4.3 mol/L [7]. In a continuous dialysis, increasing the number of AEMs can also accelerate the acid recovery apparently [40]. As an example, For instance, the number of AEMs increases from 7 to 19 if the acid recovery is increased from 35% to 78% [40]. In a batch dialysis, liquid mixing is fulfilled by stirring in both chambers [33,38]. The increase of liquid mixing intensity is proven to be favorable for DD transport in some cases, because of the elimination of the diffusive layer influence in the liquid adjacent to the membrane surfaces [33,38]. 4.2. Hydrochloric acid (HCl) recovery Since 1964, a number of steel pickling industries started to select HCl as pickling agent, because usage of HCl instead of H2 SO4 can induce faster and cleaner pickling, lower acid consumption, less quantities of waste pickle liquor and more uniform product [31]. Consequently, the research of HCl recovery by DD emerges. HCl and H2 SO4 are both strong acids, and hence the fundamental principles for their diffusion processes are similar. Nevertheless, differences of their physico-chemical characteristics can also cause some differences in their DD processes, including the “salt effect”, the effect of acid concentration on salt transport and the effect of complexes, as will be further discussed. 4.2.1. The salt effect in hydrochloric acid (HCl) recovery The “salt effect” can always be observed in the HCl system. In such a scenario, addition of salt chloride can enhance the permeability of HCl, leading to increased HCl concentration at water side and higher overall dialysis coefficient of HCl [9,17]. For instance, in the systems of HCl + FeCl2 , HCl + FeCl3 , HCl + NiCl2 , and HCl + ZnCl2 , the presences of salts enhance the permeability of HCl, so that the concentrations of HCl in water side can reach higher values after a certain running time (25–30 h) than that in the feed side [9,17,43,44]. In Section 4.1.3, the “salt effect” generally cannot be observed in H2 SO4 + sulphates systems. Reasons behind the differences are still not clearly known. The much higher ionic strength contributions of SO4 2− than Cl− may play some role. Based on a mathematic model for HCl + NiCl2 and HCl + FeCl2 DD systems, Palaty´ and Zakova [17,44] have defined four phenomenological coefficients PA–B , PA–A PB–A PB–B (A-acid, B-salt chloride) which represent the effects of acid or salt on the permeability of acid or salt. The fitted PB–A value is somewhat lower but of the same order of magnitude as the fitted PA–A value [17,44]. This finding indicates that the concentration of salt chloride is an important factor affecting the permeability of HCl, and hence can also support the existence of “salt effect” in HCl + chloride system. 4.2.2. The effect of complexes on hydrochloric acid (HCl) recovery Complexes of metal ions with Cl− can be formed and the permeability of metal ions or HCl is influenced accordingly. For the system of HCl + NiCl2 [44], the following species related with Ni2+ exist in the solution: Ni2+ and NiCl+ . Because of the high volume of NiCl+ ions and the high repulsive forces between Ni2+ and fixed charges in the membrane, the fluxes of both NiCl+ and Ni2+ through the membrane are low. Therefore, the selectivity and recovery of HCl is high [44]. Similar phenomenon is also observed in FeCl2 + HCl system [17], though the FeCl2 flux is slightly higher than that of NiCl2 . That’s because (FeCl2 )non , as well as Fe2+ and FeCl+ , is present in HCl + FeCl2 system and (FeCl2 )non can transport more easily [17]. For the system of HCl + ZnCl2 [9], Zn2+ can form a series of complexes with Cl− , i.e. ZnCl+ , ZnCl2 , ZnCl3 − and ZnCl4 2− , among which the molar fraction of ZnCl3 − is higher than 85% throughout

J. Luo et al. / Journal of Membrane Science 366 (2011) 1–16

the investigated acid concentration (0–4 mol/L), and ZnCl2 exhibits high affinity to the membrane (Neosepta-AFN). The flux of ZnCl2 is enhanced (∼1 × 10−7 kmol/m2 s, much higher than ZnSO4 flux in H2 SO4 + ZnSO4 system (2 × 10−7 kmol/m2 s)), while the separation efficiency of HCl to ZnCl2 is suppressed accordingly [9]. For the system of FeCl3 + HCl [43], components of Fe3+ , FeCl2+ , FeCl2 + and FeCl3 exist. Fe3+ shows high positive valence, while Fe complexes (FeCl2+ , FeCl2 + and FeCl3 ) have high molecular weights. Therefore, their transports are negligible and overall dialysis coefficient for HCl is high (∼1 × 10−6 m/s) [43]. 4.2.3. The effect of other factors on hydrochloric acid (HCl) recovery Besides the aforementioned factors; other factors including the HCl and salt chloride concentrations, retention time, the ratio of water to feed and rotational speed can also influence DD performances significantly. For instance, the concentrations of HCl and its salts (such as NiCl2 and ZnCl2 ) in water side increase with increase of both acid and salt concentrations in the feed side [9,44]. This can be ascribed to the increased concentration gradient across the membrane, the enhanced ionic mobility and their diffusion coefficients in the membrane, as well as the salt effect. Among these, enhanced ion mobility and higher diffusion coefficient are qualified by membrane conductivity values. It is key to DD process since mass transfer resistance mainly exists in the membrane [45,46]. As for the feed system with only HCl, the influence of feed concentration and operation conditions on DD performances can be summarized as following [47]: with increasing HCl concentration, the HCl diffusion coefficient, mass transfer coefficient and the acid uptake in the membrane increases. With increasing retention time, the HCl recovery increases linearly first, but then gradually levels off due to decreased concentration difference and increased resistances in liquid films. With increasing flow ratio of water to feed, the recovery increases, mainly due to the decreased acid concentration in water side and the increased concentration difference across the membrane [47]. With increasing intensity of liquid mixing in both sides, the diffusion coefficient of HCl increases. The trends above are similar to those of H2 SO4 system. However, the flux and diffusion coefficient of H2 SO4 are much greater than those of HCl in the low concentration range (0–0.1 mol/L) [33], because of higher acidity of H2 SO4 and higher affinity of H2 SO4 to the membrane. All of the conclusions about the feed system with only HCl can be references for understanding more complex HCl DD systems. 4.3. Nitric acid (HNO3 ) recovery A combination of hydrofluoric (HF) and HNO3 is often used for pickling agents in special metal processing industries [31], and large quantities of spent liquor are generated. For instance, 2 × 104 to 4 × 104 kg waste liquors composed of HNO3 (230–260 g/L), HF (3 g/L), Ti4+ (18–24 g/L) and other metal impurities [20], will be produced when 1000 kg titanium materials are processed. In the stripping solution for the printed circuit boards, the contents of dissociative HNO3 and Sn are more than 100 g/L (20–30%), and the contents of Cu and Fe ions are 14 g/L and 4 g/L respectively [48]. Since HNO3 and HF are more expensive than other inorganic acids, their regenerations are necessary. For the mixed acids HNO3 + HF from the titanium spent leaching solution [20], a portion of Ti4+ ions exists in complex anions form [TiF6 ]2− due to the following reactions: 12HF + 3Ti + 12HNO3 = 3TiF4 + 12NO2 + 12H2 O

(1)

TiF4 + 2HF = [TiF6 ]2− + 2H+

(2)

+

2−

Ti + 6HF = 2H + [TiF6 ]

+ 2H2

(3)

7

[TiF6 ]2− is negatively charged and more likely to leak through the membrane. The relation of [TiF6 ]2− leakage, water osmosis and acid recovery with the membrane properties, flow rate and flow ratio (water to feed) has been investigated in our laboratory, with PPObased AEM as the DD membrane [19]. The main findings are as following: both the total acid recovery and titanium leakage firstly increase and then decrease with the flow rate. And with the increase in flow ratio of water to feed, the total acid recovery and titanium leakage ratio increase. The water osmosis becomes serious when the flow ratio exceeds 1.5. Titanium leakage becomes more significant when the membrane with a higher IEC is used. In summary of the above changing trends, the flow rate 335 L/h, the flow ratio 1.05:1.1 and AEM with 16% ABC and 29% BBC are recommended for practical operation. Based on these parameters, the economic estimation is made for the HNO3 + HF recovery system of a titanium etching plant (Fig. 6.) in western China The recovery ratios for HNO3 and HF can reach 85% and 90% respectively. Such a dialyzer can generate close to CN$ 4.3 million in operational profit per year and also incur significant environmental benefits [20]. There are also some trials for HNO3 recovery from waste stripping solution of the printed circuit boards. In the work of Zhang et al. [48], HKY-001 diffusion dialyzer installed with similar PPObased AEM is used for waste solution containing 15% HNO3 (weight fraction) and 123.5 g/L Sn. The recovered HNO3 concentration is high (11.5% weight fraction), and Sn leakage is low (3.1 g/L). Another interesting point about such DD system is that a certain amount of dissociative HNO3 should be kept in the feed so as to avoid precipitation from the dissociated Sn. Hence, the HNO3 recovery rate cannot reach an optimal high value of 71% [48]. Besides empirical and analytical research, deterministic and probabilistic modeling, and schematic studies have been carried out on ionic mobilities of H+ and NO3 − in the HNO3 system dually or in comparison to the HCl system [49]. One of the differences between HNO3 and HCl DD systems is that the non-dissociated form of HNO3 needs to be considered due to a lower dissociation degree of HNO3 in the membrane. However, this leads to insignificant difference of HNO3 or HCl mobility through the membrane. The main reason can be attributed to diffusivity of the non-dissociated HNO3 (0.93 × 10−9 to 0.86 × 10−9 ) which changes imperceptibly with increasing acid concentration in the membrane [49]; and the mobility of H+ ions is also much higher than that of NO3 − and Cl− within set concentration ranges [46]. Table 2 is a summary of applications of DD systems in recovery of strong acids, as discussed in Sections 4.1–4.3.

4.4. Schematic and prototypical designs for diffusion dialysis In recent decades, U.S. and Japan have led in technological advancements of equipment associated with strong acids recovery using DD. More than fifty factories have been using DD to recover acids successfully from acidic feeds in Japan and a variety of improved DD equipment has been extended to industrial applications [1]. For instance, Exergy Technologies Corporation of United States has designed a fully automatic mode with positioning frame between films (Fig. 7(a)); Sawyer And Smith Corporation has developed diffusion dialyzer with small footprint and small stress add between the clapboard (Fig. 7(b)); Asahi Glass Co., Ltd. of Japan has developed diffusion dialyzers and IEMs suitable for dialyzers, and has made those products commercialized (Fig. 7(c)). From these configurations, the DD performance and the equipment reliability and life are improved greatly. Astom Corporation of Japan has introduced flexible rubber partition in DD equipment to improve the flow state of solution, leading to

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Table 2 Application fields and some related information (scale, membrane, characteristic and economic estimation) of diffusion dialysis for strong acids. Application field

Scale

Membrane

Process characteristic

Economic evaluation

Reference

Separation of H2 SO4 + CuSO4 mixture

Lab scale (batch dialysis)

Neosepta-AFN (Astom Corporation, Japan)

Membrane area: 62.2 cm2 ; temperature 20 ± 0.5 ◦ C; RCuSO4 > 0.965 in case of

–

[35]

H2 SO4 recovery from waste acid solution

Pilot scale: Asahi Type T-0b Dialyzer

Selemion DSV (Asahi Glass, Tokyo, Japan)

[7]

H2 SO4 and Ni recovery from electrolysis spent solution H2 SO4 recovery from waste aluminum surface processing solution

Lab scale and Pilot runs (industrial scale) Pilot scale: Asahi dialyzer (Model T-O)

Brominated PPO-based AEM (Shandong, China)

Investment cost: US$ 500,000; total savings: US$ 375, 590/year; payback period < 16 months; membrane’s life > 5 years –

[40]

Separation of H2 SO4 +ZnSO4 mixture

Lab scale (batch dialysis)

Neosepta-AFN (Astom Corporation, Japan)

Total system cost: US$ 292,500/year; total savings: US$279,600/year; payback period < 1.05 years –

[36]

H2 SO4 recovery from rare earth sulfate solution

Pilot scale: TSD-2 dialysis cell (Tokuyama Ltd.Japan) Pilot scale

DF 120-I (Shandong, China)

–

[41,50]

–

[51]

Pilot scale

DF 120-I (Shandong, China)

[52]

Lab scale

Brominated PPO-based AEM (Shandong, China)

Membrane area: 5.07 cm2 ; SH2 SO4 /Ti : 51

Sulfuric acid conservation cost: RMB 12.15 million/year –

Industrial scale

DF 120-I (Shandong, China) DF120-I (Shandong, China)

Net profit: RMB 3,112,000/year –

[54]

Pilot scale

H2 SO4 recovery from solution containing uranium H2 SO4 recovery from acid leach solution

Lab scale (cycling operation)

–

RH2 SO4 : about 80%; CH+ (recovered): 4.4–5.2 g/L Membrane area: 4 m2 ; RH2 SO4 and RNa2 SO4 : 70–80%; RCOD : about 70% Membrane area: 54 cm2 ; RH2 SO4 : 30%; RU < 10%

–

[56]

Pilot scale (HKY-001 dialyzer)

DF120-III (Shandong, China)

Membrane area: 3.2 m2 ; RH2 SO4 : 80%; RV : 93–95%; RFe : 92–94%

[32]

H2 SO4 recovery from waste anodic aluminum oxidation solution

Lab scale; Pilot scale;

DF120-I (Shandong in China)

[38]

HCl recovery

Lab scale and pilot scale (TSD-2 dialyzer)

Neosepta-AFN (Tokuyama Co., Japan)

–

[47]

HCl recovery from solution containing NiCl2

Lab scale

Neosepta-AFN (Tokuyama Co., Japan)

–

[44]

HCl recovery from solution containing HCl + FeCl2

Lab scale

Neosepta-AFN (Tokuyama Co., Japan)

Membrane area for lab scale: 2.83 cm2 ; Membrane area for pilot scale: 3.2 m2 ; H2 SO4 recovery: 85.25%; Al3+ leakage: 4.98% Membrane area: 16 cm2 for lab scale; membrane area: 200 cm2 for pilot scale; RHCl : 60–80% Membrane area: 62.2 cm2 ; permeability for HCl: 0.84 × 10−6 to 2.4 × 10−6 ms−1 ; permeability for NiCl2 : two orders of magnitude lower than that for HCl The partial flux of FeCl2 < 5.6%; the concentration gradient of FeCl2 has significant effect on the salt flux

Net benefit: 0.81 million dollars/year; payback period < 2.2 year –

–

[17]

−3

H2 SO4 recovery from titanium white acid by DD H2 SO4 recovery from hydrometallurgy leaching solution H2 SO4 recovery from titanium white (pigment) waste solution Waste acid recovery from Hua Cheng Foil Factory Acid recovery by DD in chemical fiber factory

Selemion DSV (Asahi Glass, Tokyo, Japan)

DF 120-I (Shandong, China)

CCuSO4 > 0.75 kmol m Membrane area: 0.327 m2 ; membranes number: 19; RH2 SO4 : 80–57%; RNiSO4 and RFeSO4 : 96 and 99%; SH2 SO4 /FeSO4 : 15–18, SH2 SO4 /NiSO4 : 140–270

Membrane effective area: 2.32 m2 ; RNiSO4 > 96%; RH2 SO4 : 66–72% Membrane area: 0.326 m2 ; membranes number: 19; RH2 SO4 : 82–90%; RAl2 (SO4 )3 : 35.3–38.5%

Membrane area: 62.2 cm2 ; transport of negatively charged complexes of Zn2+ being controlling step Membrane area: 0.02 m2 ; membranes number: 11; one-pass and cycling operation; RH2 SO4 : 70–80% Membrane area: 1.9 m2 ; SH2 SO4 /FeSO4 : 23.6; RH2 SO4 > 85%; RFeSO4 > 93% Membranes number: 40; RH2 SO4 : 80%; RV : 96%

[8]

[19,53]

[55]

J. Luo et al. / Journal of Membrane Science 366 (2011) 1–16

9

Table 2 (Continued) Application field

Scale

Membrane

Process characteristic

Economic evaluation

Reference

HCl recovery from the waste acid solution HCl permeability through Neosepta-AFN

Pilot scale

DF120-I (Shandong, China)

–

[18]

Pilot scale

Neosepta-AFN (Tokuyama Co., Japan)

–

[42]

HF + HNO3 recovery from titanium spent leaching solutions

Industrial scale

Brominated PPO-based AEM (Shandong, China)

RHCl > 88%; RFe : 89–77%; RZn > 56% The permeability of HCl apparently increases with acid concentration Membrane area: 512 m2 ; RH+ : 85%

[20]

Tin and HNO3 recovery from the spent solder stripper

Lab scale

DF 120-I (Shandong, China)

Profits: CN$ 0.86 million/year; payback period < 5 months –

an increased acid recovery (Fig. 7(d)). DD and allied technological development has also displayed remarkable growth in recent years in China. For instance, Shandong Tianwei Membrane Technology Co., Ltd owns independent intellectual property rights of homogeneous membrane production, and has equipped 20 diffusion dialyzers with a total membrane area of 10,240 m2 (Fig. 7(e)) mainly for HCl and H2 SO4 recovery from Aluminum Foil production feeds. Economic benefits can be as high as RMB (Chinese Dollars) 1,000,000/year and nearly 59,000 m3 /year feed acidic solutions can be treated, or 9,000,000 kg/year initial acids recovered [1]. We have also invented a new compact spiral module for DD with an attractively reduced total equipment weight (Fig. 8) [57]. One piece of IEM and one mesh fabric are rolled with two pipes to form the main part of the module and two flow channels. The rolled IEM is then sealed at the edges. Thereafter, two pipes are axisymmetrically fixed along the edges of the IEM to form additional pair of flow channels. Subsequently, the formed cylinder is wrapped tightly with plastic film and placed in a columnar hard surface. The two ends of the hard surface are sealed with adhesive to form the final module with two inlets and two outlets as shown in (Fig. 8(a)). The main characterizations of the module are its compact nature, high filling density and small foot print. During the DD process, spirally radial counter current across the membrane and high turbulent degree can be realized, resulting in high mass transfer and increased feed treatment efficiency. Besides, the spiral module can also be easily integrated with other reaction or separation equipment due to its small size and simple operation [57].

Membrane area: 95 cm2 ; RHNO3 > 70%

[48]

5. Treatment of organic acid using diffusion dialysis Organic acids have been widely used in chemical, leather, food, fermentation and pharmaceutical industries. Among them, the bacterial fermentation industries always generate carboxylic acids and carboxylates of various compositions and concentrations [58]. Acids are toxic for bacterial during fermentation process, and the pH of the broth should be in the range of 5–6 [58,59]. Hence different methods [59,60], including facilitated membrane extraction, neutralization dialysis, ion exchange resin, integrated systems of ultrafiltration and electrodialysis have been used to remove acids from the broth. Meanwhile, DD, as a more simple and economical method, has been studied for the recovery of weak acids from those industries. We further discuss some studies relating to DD recovery of carboxylic acids. 5.1. The factors determining the separation of weak acid from its salt Similar to strong acid, the separation of weak acid from its salt is determined by their sorption and diffusion differences. Nevertheless, the sorption and diffusion properties of weak acids follow different rules as compared with strong acids. Low absorption and high diffusivity of strong acids within membrane matrices are always observed. For instance, the acid concentrations of H2 SO4 , HCl and HNO3 in membrane are always lower than that in feed solution [35,47,49], and the diffusion coefficients (D) for HNO3 and HCl are ∼2 × 10−6 to 6 × 10−6 m2 /s, 2–5

Fig. 6. Mixed acid (HNO3 + HF) recovery flow sheet from titanium material processing industries using diffusion dialysis. Tl, recovered acids collection tank; T2, fresh acids tank; T3, etching tank; T4, spent waste liquor tank; T5, spent waste liquor high-position tank; T6, stripping water high position tank; T7, diffusate waste collection tank; C1, neutralization cell; C2, precipitation cell [20].

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Fig. 7. Some representative diffusion dialyzers in USA, Japan and China.

orders of magnitude larger than that of their salts in membrane [47,49]. Therefore, the successful separation of strong acid from its salt is mainly due to their high diffusion rates. In contrast, the sorption of weak acids can be much more significant inasmuch as their diffusivities are relatively low. For the

system of lactic acid (LA) + NaLA, the LA concentration in the membrane is equal to or even higher than that in external solution [58]. The concentration of acetic acid (AA) or propionic acid (PR) in the membrane is higher and in some cases can be several times higher than that in external solution [58]. Especially, when the PR con-

Fig. 8. Prototype DD configuration (a) schematic and (b) image of the spiral module for diffusion dialysis recently developed in our facilities [57].

J. Luo et al. / Journal of Membrane Science 366 (2011) 1–16 Table 3 Values of molar mass (MA ) and dissociation constant (Ka) for different carboxylic acids. Acid

MA (g/mol)

Ka

pKa

Acetic acid (AA) Oxalic acid (OA) Lactic acid (LA) Tartaric acid (TA) Citric acid (CA)

60.04 90.03 90.08 150.09 192.13

1.76 × 10−5 6.5 × 10−2 1.4 × 10−4 9.21 × 10−4 8.6 × 10−4

4.75 1.23 3.86 3.04 3.13

centration in external solution is within 0.75–1.0 mol/L range, its concentration in the Neosepta AFN membrane is roughly four times higher (2.6–3.6 mol/L). This high sorption can be possibly attributed to be the association of the acids with the water molecules incorporated in associates [58] even though the diffusion of weak acid is relatively low. For instance, the D values for LA and its salt are ∼10−11 m2 /s [59]. Correspondingly, it can be deduced that a high sorption of weak acid determines the effective separation of weak acid from its salt [59]. Some of the separation results are: the flux of the LA is up to 1 mol/m2 h, 10 times higher than the flux of the salt; the fluxes are as high as 1.9–2.0 mol/m2 h for AA and 1.5–1.9 mol/m2 h for PR, while the fluxes of their salts do not exceed 0.08 mol/m2 h [58,59]. Due to the importance of sorption for weak acids separations, for sorption studies of different weak acids the following conclusions can be drawn [61]: the acid concentration in membrane sharply increases with increasing acid concentration at low acid concentration, but decreases with acid concentration increasing further. The sorption of acetic acid (AA), lactic acid (LA), oxalic acid (OA) and tartaric acid (TA) at room temperature follows the order of LA < AA < OA < TA. Molar mass and dissociation constant are the main determining parameters (Table 3.) and the interaction of weak acids with the membrane matrix can also play some role [61]. 5.2. Transport characteristics (diffusivity, permeability, mass transfer coefficient and overall dialysis coefficients) Diffusivity (D) of acid depends on its concentration in membrane and its interaction with membrane matrix. The D values of AA, LA, OA and TA in the Neosepta-AMH membrane are very low and approximately two orders of magnitude lower than that in aqueous solution (10−9 m2 /s) [61]. Different changing trends of D

11

values for AA, LA, OA and TA can be observed with increasing acid concentration in the membrane in the order DTA < DLA < DAA < DOA , which is similar to that for aqueous solution [61]. Permeability of membrane (P) is related with the acid concentration and its diffusivity in membrane. Generally, the P values of weak acids decrease with increasing acid concentration and increase with increasing temperature [15,34]. The temperature impact can be ascribed to increased diffusivity of the acid inside the membrane [15,34]. As for the above four acid systems, the P values follow the orders: PLA < PTA < PAA < POA , and POA is one order magnitude of larger than that of others. The membrane mass-transfer coefficient is influenced by D or P comparatively [34]. The membrane mass-transfer coefficient for OA is about one order of magnitude higher than that for citric acid (CA) or TA. And the values of AA, LA, OA and TA also follow different changing trends [34]. Overall dialysis coefficient decreases with increasing acid concentration in feed. Among LA, TA, AA and OA, the coefficient of OA is about one order of magnitude larger. Besides, higher rotational speed of the stirrers is found to be favorable for the overall dialysis coefficient in the case of OA [34]. Various results from applicability of DD systems for recovering weak acids are tabulated in Table 4. 6. Base recovery using diffusion dialysis 6.1. Alkali waste and treatment Alkali waste is mainly generated from paper, leather, printing and dying, tungsten ore smelting, and man-made fiber industries [3,6,13,24,63]. The direct discharge of the waste would lead to: (1) corrosion of channels, pipelines of plants; (2) change of water pH, affecting the self purification of rivers and other water bodies; (3) death of fishes due to the depletion of dissolved oxygen by large amount of organics in alkaline wastewater; (4) metabolic disorders in humans. Therefore, different methods are adapted to deal with alkali feeds [3,6,13,24,63], such as neutralization with acids, concentration and burning (in paper industry) and many membrane-related methods (electrodialysis, reverse osmosis, IEM electrolysis, diffusion dialysis, and ultrafiltration/nanofiltration). The neutralization method always consumes large quantities of acids and generates

Table 4 Application fields of diffusion dialysis in organic acid separation. Research and application field

Scale

Membrane

Process characteristic

Reference

Transport of acetic acid

Lab scale

Neosepta AMH (Astom Corporation, Japan)

[62]

Separation of lactic acid/lactates from fermentation products Separation of carboxylic acids from carboxylates

Lab scale

Neosepta AMH (Astom Corporation); Selemion DSV (Japan) Neosepta AFN (Astom Corporation); Selemion DSV (Japan) Neosepta AMH (Astom Corporation, Japan)

Membrane area: 100 cm2 ; temperature: 50 ◦ C; acid flux is proportional to acid concentration, membrane permeability is independent of acid concentration Membrane area: 50 cm2 ; fluxes: 1 mol/m2 h for acid, 0.07 mol/m2 h for salt separation factor: 20 for Neosepta AFN, 30 for DSV membrane Separation factor: 20–37 for AFN, about 29 for Selemion DSV; Fluxes: 1.5–2.0 for acid, 0.07–0.08 for salt Membrane area: 62.2 cm2 ; the mass-transfer data of oxalic acid is the largest (1.0 × 10−6 to 1.0 × 10−7 m/s, nearly one order of magnitude higher than others) Membrane area: 3.31 × 10−2 m2 ; P values of AMH: 7.50 × 10−9 to 3.57 × 10−7 m/s, and its order: PCA < PLA < PTA < PAA < POA Membrane area: 7.07 cm2 ; AFN membrane shows the highest mass-transfer coefficient The partition coefficients (acid concentration in membrane to that in external solution) decrease with increasing acid in feed; The orders of D values: DTA < DLA < DAA < DOA

Pilot scale

Transport of some carboxylic acids

Lab scale (batch dialysis)

Continuous dialysis of carboxylic acids

Pilot scale

Neosepta AMH (Astom Corporation, Japan)

Transport of formic acid

Lab scale (batch dialysis)

Solubility and diffusivity of carboxylic acids in membrane

Pilot scale

SB-6407 (Gelman Sciences); Neosepta AMH, AFN, ACM (Japan) Neosepta AMH (Astom Corporation, Japan)

[59]

[58]

[34]

[15]

[60] [61]

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large amount of sludge [63]. The concentration and burning method is mainly used for treatment of paper industry waste. The requirements for the processing procedures and the quality of waste feed are rigorous and the energy consumption is high. Hence this method cannot be well utilized in most developing countries. For instance, there are no economical and advisable methods to deal with the waste that is generated 1 × 106 to 3 × 106 kg/year in paper industries in China [24]. Among the different membrane-related methods, the energy cost of DD is the lowest, although the processing efficiency is relatively lower. Besides, DD is more desirable under some conditions. For instance, when it is difficult to add external driving force to treatment systems, or concentration difference between feed and water is relatively high, DD can play its unique role. In the following section, the principle, application and development of DD in alkali recovery will be assessed [63]. 6.2. The application and development of diffusion dialysis for alkali separation DD technique has been mainly used for acids recovery. Until around 20 years ago, Astom Corporation of Japan developed alkali-resistant CEM, which was a remarkable progress in DD [3]. Furthermore, Astom Corporation also successfully developed a DD process to recover NaOH from the aluminum etching solution. These earliest trials and experiences make significant contribution to the realization of alkali recovery through DD. The production technique was industrialized in a California Caspian plant, USA, around 1991 incorporating diffusion dialyzer TSD10-300 and TSD25-250 [3]. The following reaction can occur during the DD process of aluminum etching solution: NaAlO2 + 2H2 O = Al(OH)3 + NaOH

(4)

The rate of reaction (4) is slow. Accordingly, a schematic production flow is depicted as Fig. 9: alkali is recovered through DD first, then the deplete solution is transported to a crystallizer outside the dialyzer. In the crystallizer, Al(OH)3 is formed and recycled. From Fig. 9, membrane fouling due to precipitating Al(OH)3 can be eliminated. The recovered NaOH concentration can be even slightly higher than that of the feed, since the driving force for Na+ migration is high, provided by the Na+ ions from both NaAlO2 and NaOH. Economical estimation indicates that the annual net profit is 70,000 dollars, and the investment recovery period is about 2 years [3]. Inspired by initial successes of DD, recovery of alkali from the alkali etching solution of white tungsten ore smelting industry has been tried with dialyzer TSD-2-20 in Japan [11]. The feed alkali solution contains WO3 91.25 g/L and NaOH 1.36 mol/L. When the flow rate of the feed and the recovered solution are controlled at 40 mol/min and 10 mol/min respectively in a cyclic manner, the concentration of recovered solution is up to 51% of feed concentration, and the recovery rate is 50% [11]. Variation of the operation conditions demonstrates that the migration speeds of NaOH and tungsten follows different rules: the migration speed of NaOH increases with increasing feed flow rate and feed concentration, but remains almost unchanged with the water flow rate. Meanwhile, the migration rate of tungsten changes insignificantly with all the three variables (the flow rates of feed and water, and feed concentration). Therefore, the speed ratio of NaOH to tungsten can be up to 600–700 with the increase of feed flow rate, and up to 2000 or even higher with an increase of feed concentration [13]. Another finding is that the volume of feed increases apparently during the DD process because of water osmosis. Hence water osmosis is a key factor to affect the DD performance and should not be ignored in the alkali recovery. In paper industry, grass fiber and wood fiber are always used as raw materials, so high contents of cellulose, lignin, alkali (pH

Fig. 9. The process for NaOH recovery from aluminum etching solution through the combination of diffusion dialyzer and crystallizer [3].

11–13) as well as high COD (Chemical Oxygen Demand) content always exist in the feed solution, making the waste treatment complex and difficult. Macromolecule Institute of Shanghai Jiao Tong University of China selected three kinds of membranes (IEM from polyethylene, hydrophobic and hydrophilic fluoride membranes) to recover NaOH from alkaline feed solution [24]. The diffusion dialyzer is divided into three compartments by an AEM and a CEM (Fig. 10). The OH− and Na+ pass the AEM and CEM membranes into water respectively, and then water in the two diffusate chambers are mixed together through a pump. The hydrophilic fluoride membrane performs the most outstanding function: the alkali recovery efficiency and rejection for organic substance are higher and running/service life is longer than the other membranes [24]. The above research and application results show that the DD technique has wide potential for the alkali separation. However, the recovery ratio and separation effect is not high enough in some cases, due to the native characteristics of the alkali solution (the paper industry waste, for instance) and the existence of high water osmosis for CEM. Therefore, the alkali recovery with DD is often utilized as the pre-treatment step for electrodialysis. The burden of alkali recovery can be reduced and the recovered dilute alkali solution can be used as electrolyte for cathode chamber. 7. Integrated processes based on diffusion dialysis As discussed in the introduction section, the integrations of DD with other methods are important means to overcome the disadvantages of DD and to get more excellent treatment results. The

Na + OHFeed

1.Peristalticpump 2.Anionexchangemembrane 3.Cationexchangemembrane 4.Aeratorpipe

Fig. 10. Diffusion dialysis device by Macromolecule Institute of Shanghai Jiao Tong University (China) for alkali treatment [24].

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shortcomings of the DD process are mainly its low processing capability and efficiency. For instance, DD cannot recover acids from feed completely and then the neutralization of the depleted solution with alkali results in the loss of metal salts. DD cannot reclaim the values of the metal effectively in the acidic feed solution. DD is not effective for feed with low concentration or feed with solid particles (membrane block, pollution and other undesirable effects may occur).

13

extractants are costly. Besides, the recovery of nuclides and the regeneration of the extractants are difficult. Therefore, other more effective methods are needed. Mathur et al. [67] have investigated the utilization of electrodialysis for nuclides removals. They found that the nuclides can transfer from the feed compartment into the electrodialysis cell effectively only when the acid concentration is low (0.4 mol/L). Accordingly, DD method is used as a pretreatment step and the HNO3 content in the waste is reduced to 0.3 mol/L through DD [67]. Separation rate of Cs+ and Sr2+ ions by subsequent electrodialysis is found to be between 98 and 99% [67].

7.1. Diffusion dialysis integrated with electrodialysis 7.2. Diffusion dialysis integrated with IEM-electrowinning The depleted solution from DD is always neutralized with alkalis and some metal salts are generated accordingly. In order to recover the metal values, electrodialysis, which is a electro-driven membrane process and has been widely utilized for the ion concentration, removal and enrichment [64], can be adopted. As shown in Fig. 11, Negro et al. [10] presented an integrated treatment method for treatment of HNO3 and HF solution. DD is the first step to remove part of the acids (recovery rate: 97% for HNO3 and 50% for HF). Then the depleted waste solution is neutralized with KOH. The formed metal sludge is separated and drained with a filter press. Subsequently the separated salt solution is split with high efficiency in a bipolar membrane electrodialysis. The produced acid is recycled to the pickling bath and the base to the precipitation process. The depleted salt solution from the electrodialysis process is sent to a standard electrodialyzer with monopolar membranes to upgrade its concentration. A closed loop of water can be achieved by the reuse of the spent acid stream from DD as input in the acid compartment of the salt splitting unit, then by the reuse of the dilute stream from the standard electrodialysis unit as the water stream for DD. Nearly no solid waste is produced. Furthermore, the recovery of valuable chemicals can contribute to the economical viability of the process. The payback time of the proposed process is less than 2.5 years [10]. The high-level acid (∼3.0 mol/L) wastes with radioactive nuclides 137 Cs and 90 Sr are generated in the production process of nuclear fuels [65,66]. Different kinds of extractants have been used to remove those nuclides, such as crown ether and di-tert-butylcyclohexano 18-crown-6 (for 90 Sr), ammoniummolybdophosphate and hexacyanoferrates (for 137 Cs) [65,66]. Nevertheless, these

The spent solutions from the printed circuit board industries contain high content of acids and other metals [48]. The Sn content can be up to 100 g/L, the Cu and Fe contents are nearly 20 g/L, and the acid content is ranged from 20 to 30% in some cases. For recovery of the valuable chemicals, the DD and IEM-electrowinning integrated process is developed as a new alternative method [68]. Firstly, HNO3 is recovered by DD, and then the deplete solution undergoes the IEM-electrowinning process. The standard electrode potential difference of Cu and Sn is greater than 0.2 V, hence they can be recovered by controlling the cathode potential of the IEMelectrowinning process. The results show that the HNO3 recovery is up to 70%, and Sn recovery is up to 62% with current efficiency up to 60% [48,68]. 7.3. Diffusion dialysis integrated with continuous membrane extraction Continuous membrane extraction (CME) is widely applied for the recovery and concentration of different metal ions. It is especially useful for the separation of the metal ions with the same valence. The setting for CME is quite simple; i.e. the feed and strip aqueous solutions are separated by two IEMs, between which an organic solution containing extractant flows. The solutes in the feed transport through the membrane and are then extracted into the organic phase. Subsequent stripping process occurs between the organic solution and the strip solution [69]. In the work of Wodzki et al. [69], CME is integrated with DD to separate Cu2+ and Zn2+ from the electroplating waste. Cu2+ and

Fig. 11. The integrated process of diffusion dialysis and electrodialysis for pickling waste recovery.

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Fig. 12. The integrated process of vacuum membrane distillation and diffusion dialysis for acid recovery.

Zn2+ are firstly separated from the waste in a CME, with H2 SO4 as the stripping medium. Then the products of the CME process are further de-acidified by DD. In DD process, the selectivity of acid to salt attains the high value of 166. With the integrated DD and CME processes, the flux of Cu2+ and Zn2+ reaches 97.1% of total ion flux [70].

brane is used for TiO2 clarification [73]. After microfiltration, TiO2 particles can be recycled and the turbidity of the waste is lowered to less than 0.5 NTU (NTU denotes turbidity). Subsequent DD process is reliable. The H2 SO4 recovery rate is more than 85% and the leakage rate of salt is less than 7% [74]. 8. Summary and perspective

7.4. Diffusion dialysis integrated with vacuum membrane distillation DD is always effective for the feed solution with high concentration. In the case of low acid waste, the investment is increased. Furthermore, the recovered acid has low concentration and cannot be directly re-used [8]. Hence some other processes can be utilized for pre-concentration, such as vacuum membrane distillation (VMD) [71] (Fig. 12). VMD is a low temperature and vapor–pressure driven process and has been widely studied for the concentration of aqueous solution with non-volatile solute, etc. [71,72]. In the rare earth smelting process, large amount of acidic feed is generated with low acid concentration (<0.5 mol/L). Hence the feed is firstly concentrated through VMD, followed by treatment with DD [71]. The investment is decreased, and the recovered acid concentration and the processing speed are increased. H2 SO4 in rare earth sulphate solution can be also concentrated and the recovery rate is between 70 and 80% [50,70,71]. 7.5. Diffusion dialysis integrated with ceramic membrane microfiltration During DD process in some cases, the waste solution containing some solid wastes can lead to the block and pollution of the membrane, the decrease of the treatment efficiency and membrane service life. Accordingly, the waste solution should be purified before DD process. The ceramic membrane microfiltration is advisable for solution purification due to their inherent merits of physical integrity, chemical resistance and thermal stability [30]. The acidic waste from TiO2 plant generally contains 1–2% hydrated TiO2 suspension besides acids and metal ions (20% H2 SO4 , 8% FeSO4 , etc.). The direct treatment of the waste with DD process is difficult since TiO2 particles can cause serious AEM fouling and influence the DD efficiency. Thus a ceramic microfiltration mem-

Diffusion dialysis (DD) provides an alternative, compatible and prospective method for acid and alkali recovery. The most outstanding advantages of DD are the clean nature, low consumption of energy, low installation and operating cost. Besides, the integrations of DD (particularly for acid recovery systems) with other methods can increase the processing capability and efficiency, so that the acids (even with low concentration) and the metal salts can be successfully recovered. Various parameters will influence the efficiency of DD, such as membrane properties (especially water content (WR ) and fixed group (CR )), variety and concentration of waste solution, and the operation conditions (operation temperature, the flow rate, flow ratio). Accordingly, significant differences have been reported for different DD systems: the “salt effect” is obvious in the HCl + chlorides DD systems, while in the H2 SO4 + sulphates DD systems, the “salt effect” nearly cannot be observed; low absorption and high diffusivity in membrane are always observed in the strong acid solutions, while in the weak acid solution, the trends are just opposite; the “water osmosis” is relatively more obvious in the alkali feed solution than that in the strong acid feed, and hence the efficiency of alkali recovery is lower. Especially in America, Japan as well as China, the application and development of DD has attained great achievements, producing excellent economic and environmental benefits. Nevertheless, further improvement is needed. For example, the capability and efficiency of DD, especially for weak acids and alkali recovery, is relatively low as compared with electro-dialysis and other electricdriven or pressure-driven membrane processes; the variety of DD membranes is limited due to the high requirement for physicochemical stabilities, including the stability at high temperature and in acid/alkali solution. Hence, the development of DD membranes for DD running under more rigorous conditions is highly desirable. For instance, when DD is operated at high temperature, there is

J. Luo et al. / Journal of Membrane Science 366 (2011) 1–16

no need to cool down the industrial effluents. Furthermore, the protons or hydroxyl ions will transport faster at elevated temperature, while the transportation of the salts remain approximately unchanged. Thus high fluxes of acids and alkali as well as high selectivity can be attained. Acknowledgements This project was supported in part by The Natural Science Foundation of China (Nos. 21025626, 20974106 and 20636050), The National Basic Research Program of China (973 program, No. 2009CB623403) and The Significant and Key Foundations of Educational Committee of Anhui Province (Nos. KJ2010A330, ZD2009001, KJ2009A003 and KJ2010A265).

Nomenclature IEM ion exchange membrane AEM anion exchange membrane CEM cation exchange membrane PPO poly(2,6-dimethyl-1,4-phenylene oxide) BBC benzyl bromine content ABC aryl bromine content WR water content IEC ion exchange capacity (mol/g dry membrane) CR fixed group concentration (mol/L) concentration of A (mol/L) CA Racid recovery rate of acid Rmetal salts rejection rate of metal salt Sacid/salts separation factor of acid to salt rejection rate of COD RCOD D diffusivity of acid (m2 /s) P permeability of membrane (m/s) the permeability of B affected by A (m/s) PA–B

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