Lithium-ion battery recycling through an integrated electro-membrane crystallization technology

Domain: chemistry
Published: 2025-12-20
Editorial outcome: major_revision

Abstract
Lithium-ion battery (LIB) recycling is crucial for energy security, environmental sustainability, and economic viability, as the finite lifespan of LIBs results in a significant annual accumulation of spent units. However, effectively and precisely recovering valuable metal ions such as Li+, Mn2+, Ni2+ and Co2+ from complex LIB leaching solutions remains a major challenge. Here, we present a scalable electro-membrane crystallization-assisted general recycling (e-MCGR) technology for the selective and efficient recovery of those metal ions from LIB leaching solutions. By synergistically integrating electrochemical and electro-membrane technologies, our proposed technology incorporates four key configurations: selective membrane dual-stage distillation, bipolar membrane in-situ crystallization, membrane metal-complexing ex-situ crystallization, and membrane metal-extracting temporal crystallization. We systematically analyze the metal-ion transfer kinetics of electro-membranes (ion selectivity and permeation rates) and the performance of electro-membrane crystallization systems (recovery rates and product purity), alongside evaluating energy consumption, economic viability, and environmental benefits. Our optimized e-MCGR process achieves a recovery of 95.5% for Li+, 99.5% for Mn2+, 83.1% for Ni2+, and 87.3% for Co2+, yielding high product purities of 99.9% for Li2CO3, 99.9% for Mn3O4, 99.5% for Ni(OH)2 and 92.5% for Co(OH)2. The e-MCGR technology demonstrates significant potential to reshape LIB recycling paradigms, aligning technological efficiency with energy, environmental and economic sustainability.

Introduction
Lithium-ion batteries (LIBs), the powerhouse of modern electronics, have been a global cornerstone to power billions of electronic devices1. They have undergone exponential and fast development in adding resilience to electrical grids with increasing contributions from intermittent sources2,3. By 2024, global LIB production reached approximately 3950 gigawatt-hours (GWh), with China contributing over 2677 GWh of this output, which represented over 65% of the global market share (Fig. 1a). Such rapid expansion of the LIB production underscores China’s success in reducing reliance on carbon-based energy while advancing renewable energy adoption (Supplementary Fig. 1). Unfortunately, the finite lifespan of these batteries (typically 2–9 years, around 500–3000 cycles) results in a substantial annual accumulation of spent batteries4. According to the data, the global generation of spent LIBs surged 1700 Kilotons (Kt) in 2024, with the volumes expected to rise further as demand grows (Supplementary Fig. 3). This escalating influx of spent batteries poses critical energy, economic, and environmental challenges, particularly in major economies with high LIB usage, such as China, the United States, European nations, Japan, and Korea (Fig. 1b and Supplementary Fig. 2).

The efficient recycling of spent batteries has become an imperative priority5–7. Among the complex composition of spent LIBs, the recovery of high-value metal ions, such as Li+, Mn2+, Ni2+, and Co2+, offers a potential opportunity for resource reclamation8,9. Typically, spent LIBs are dissolved in acidic solutions, followed by pH adjustment (~7) to obtain a leaching solution10, from which metal ions need to be recovered efficiently and precisely. To address these challenges, various approaches such as adsorption, ion exchange, solvent extraction, precipitation, or evaporative crystallization have been explored11–14. While some of these methods enable metal ion recovery, their limitations stem from imprecise separation and inefficient purification15,16. Thus, achieving metal ion recovery with both high recovery and high-purity products in a sustainable manner remains a significant hurdle. Additionally, most existing approaches focus narrowly on technical aspects, often overlooking system-level integration and sustainability17–19.

In recent decades, membrane technologies with tunable structures and functional groups have shown great promise as an efficient method for ion separation20–22. Among them, electrodialysis, by using electro-membranes, has emerged as a viable approach for recovering ion resources from diverse streams23–25. Advances in material science have led to the development of electro-membranes with tailored ion selectivity, enabling precise targeting of specific ions even in the presence of competing species with similar physicochemical properties26–30. However, ion transport across the electro-membranes is constrained by their ion-transfer kinetics (e-MK), resulting in trade-offs between ion selectivity and permeation rates13,31–33. Notably, recent research efforts have been directed toward combining electrochemical systems with crystallization to enhance metal ion recovery, resulting in the development of electrochemical crystallization. This approach enables simultaneous metal ion separation and crystal product generation24,34–36. Yet, a trade-off between the high ion recovery rate and product purity exists at the process scale37–39. To address this, hybridizing electrochemical systems with electro-membrane technology and crystallization technology (e-MC) offers a promising pathway for achieving metal ions with both high recoveries and high product purities. Previously, membrane technologies have relied on principles such as size exclusion, electrostatic interactions, and chemical affinity to achieve ion selectivity31,40–42. Although substantial progress has been made toward developing membranes with precise single-ion selectivity, their utility in real-world resource recovery remains limited by material constraints and has largely been tested only in isolated/single systems43–45. Furthermore, their applicability in LIB recycling has not been convincingly demonstrated42,46.

To bridge this gap, we propose an e-MC-assisted general recycling (e-MCGR) concept, which integrates mature membranes into a closed-loop and modular framework tailored for spent LIB recycling. The e-MCGR concept is designed to address both separation precision and practical applicability while maintaining high recovery and product purity. As illustrated in Fig. 2, the process begins with disassembly and acidic dissolution of spent LIBs, followed by pH adjustment (~7) to generate a representative leaching solution47. Valuable metals ions, such as Li+, Mn2+, Ni2+ and Co2+, are then selectively recovered through a three-part process: (i) separation of monovalent cations (Li+ and Na+) from divalent cations (Mn2+, Ni2+ and Co2+) through selective membrane dual-stage distillation; (ii), separation of monovalent cations from each other (separation Li+ and Na+) and Li+ crystallization via bipolar membrane in situ crystallization and (iii) selective recovery of Mn2+, Ni2+ and Co2+ by membrane metal-complexing ex situ crystallization and membrane metal-extracting temporal crystallization. These processes enable the recovery of high-purity Li2CO3, Mn3O4, Ni(OH)2 and Co(OH)2. Despite the strategic importance of metal ion recovery from spent LIBs, no comprehensive study or design has yet explored a fully sustainable LIB recycling through this e-MCGR concept. Furthermore, the potential of LIB recycling to enhance energy security, environmental sustainability, and economic viability remains insufficiently explored. This e-MCGR presents an innovative framework and practical strategy for optimizing metal ion recovery, thereby addressing the challenges encountered in sustainable resource recovery from LIBs. Given the global scale of LIB development, this e-MCGR concept and approach contribute a holistic, scalable, and sustainable solution to address global challenges in LIB waste management and resource recovery.

Results
Separation of monovalent cations from divalent cations
The e-MCGR starts with ultrafiltration (UF, Supplementary Table 6), where the organic pollutants in the feed are removed to minimize membrane clogging. Then the metal ions in the feed are concentrated using reverse osmosis (RO, Supplementary Table 6) before entering the ED cell due to low metal ion concentration in spent LIB leaching solutions (Fig. 2, Supplementary Table 7). Due to the unique thermodynamic and physical characteristics of metal ions, such as specific hydrated diameters and hydration energy (Supplementary Table 8), monovalent metal ions (Li+ and Na+) can be separated from multivalent metal ions (Mn2+, Ni2+, and Co2+). Recent advancements have enabled the use of commercial monovalent ion-selective electro-membranes, which operate based on mechanisms such as size exclusion and electrostatic repulsion (Fig. 3a and Supplementary Fig. 5). To efficiently recover those valuable metal ions, Li+ and Na+ are separated from Mn2+, Ni2+, and Co2+ through continuous selective electrodialysis. The separation is firstly evidenced by the change in color of the leaching solution (Fig. 3b), where the dark brown of the leaching solution is changed to a dark red solution containing Mn2+, Ni2+, and Co2+, while the solution containing only Li+ and Na+ is colorless.

To enhance separation efficiency, a multi-stage selective electrodialysis could be employed, but it is limited by practical constraints in metal ion recovery rates and the trade-offs in energy, economic, and environmental benefits48. Therefore, a more suitable two-stage selective electrodialysis configuration is designed here, termed as selective membrane dual-stage distillation (Supplementary Fig. 6). The impact of current density on e-MK is evaluated by using three current densities (i.e., 1.0, 2.5, and 5.0 mA cm−2). In the first-stage selective electrodialysis at operating condition of 1.0 mA cm−2 with 4 L of leaching solutions in the feed chamber, the concentrations of Li+ and Na+ increase in the recovery chamber (Fig. 3c). In this configuration, multivalent metal ions (Mn2+, Ni2+, and Co2+) are rejected by the selective membranes, while Li+ and Na+ readily permeate (Supplementary Fig. 7). However, residual Mn2+ (0.06 g L−1), Ni2+ (0.06 g L−1) and Co2+ (0.07 g L−1) mixed with amounts of Na+ (0.04 g L−1) and Li+ (0.14 g L−1) in transitional chambers necessitate further separation due to the limited selectivity of commercial membranes. Thus, the 2 L transitional solutions (in transitional chamber-1 and transitional chamber-2) from the first-stage selective electrodialysis are fed into the second-stage selective electrodialysis for additional separation (Fig. 3d). This selective membrane dual-stage distillation configuration proves optimal for separating Li+ and Na+ from Mn2+, Ni2+, and Co2+. With the combination of first-stage operating at 1 mA cm−2 and second-stage at 0.5 mA cm−2, this dual-stage distillation achieves the highest accumulative recoveries of 96.5% for Li+ and 96.4% for Na+, while completely rejecting Mn2+, Ni2+, and Co2+ (Fig. 3e and Supplementary Fig. 8).

Effectively, the transport of metal ions from the feed chamber to the recovery chamber is governed by e-MK and reflected by voltage change (Supplementary Fig. 9). Regardless of stages or operating conditions, the e-MK is the critical factor for determining metal ion selectivity, permeation rate, and ultimately, recovery and purity. Although additional stages could enhance metal ion recovery purity, they lead to higher losses of Li+ and Na+ in the transitional chamber and higher energy consumption. Thus, the selective membrane dual-stage distillation technology strikes an optimal balance for separating monovalent cations (Li+ and Na+) and multivalent cations (Mn2+, Ni2+, and Co2+) with current membranes technologies.

A comprehensive evaluation of energy consumption, economic profitability, and environmental benefits is conducted for this selective membrane dual-stage distillation configuration (Supplementary Fig. 10). With this optimized configuration, most of this energy is consumed by metal ion transport in membranes. For this selective membrane dual-stage distillation, key metrics, such as energy consumption (EC), energy cost (Ce), investment cost (Ci), Li+ recovery cost (CLi), monovalent metal ion recovery rate (Rm), and purity of recovered Li+ (PLi), are analyzed and summarized (Supplementary Fig. 10 and Supplementary Table 9). For the first-stage selective electrodialysis, the impact of current density on separation kinetics is evaluated at 1.0, 2.5, and 5.0 mA cm−2. At 1.0 mA cm−2, the system demonstrated the lowest energy cost (1.24 $ kg−1), investment cost (3.78 $ kg−1), and the energy consumption (11.83 kWh kg−1), along with the highest monovalent cation recovery rate (87.7%), compared to 66.8% and 60.5% at 2.5 and 5.0 mA cm−2, respectively (Fig. 3f). To further enhance Li+ recovery, transitional solutions from the first-stage selective electrodialysis are processed in a second-stage selective electrodialysis operated at a lower current density (0.5 mA cm−2). Based on the operating conditions and the expenditure per 1 kg of LiCl recovery, this configuration achieves around 96.5% recovery of monovalent metal ions (Li+ and Na+), with an energy cost of 2.73 $ kg−1, investment cost of 13.25 $ kg−1, and energy consumption of 15.97 kWh kg−1 (Fig. 3g). This initial separation of monovalent and multivalent cations is essential for efficient resource recovery, both technologically and in terms of energy, economic, and environmental reasons. While selective membrane dual-stage distillation technology achieves a high separation efficiency, current commercial membranes are limited by their metal ion throughput. In addition, they suffer from severe membrane scaling (Supplementary Figs. 11 and 12). On the one hand, this scaling can be attributed to the inherent poor anti-scaling properties of commercial electro-membrane. On the other hand, water splitting may occur on the CIMS surface as most divalent ions are electrostatically repelled and cannot serve as current carriers. The resulting shortage of charge carriers induces forced dissociation of water into H⁺ and OH⁻, with OH⁻ reacting with accumulated metal ions to form insoluble hydroxide precursors that deposit on the membrane surface (Supplementary Fig. 13).

Recovery of Li+ and Na+
Following the initial separation configuration, the recovered feed contains only monovalent metal ions (Li+ and Na+), yet the lack of highly Li+-selective electro-membranes poses a significant challenge, and current electrodialysis cannot achieve crystallization. Our innovative approach offers a solution that enables simultaneous metal ion separation and crystallization of Li2CO3. Typically, CO2 is captured by reacting with OH−, a method commonly employed for carbon capture and storage, and utilized as a carbonate source. Leveraging the bipolar membrane electrodialysis, an in situ crystallization configuration is developed to separate Li+ from Na+ while generating high-purity Li2CO3 crystals (Supplementary Fig. 14). In this system, CO2 is captured via a membrane contactor (Supplementary Fig. 15) and reacts with the OH− generated in situ by the bipolar membrane (chemical reaction pathways are detailed in Supplementary Fig. 16). This hybridization of bipolar membrane and in situ crystallization technologies ensures Li+ capture, CO2 utilization, and Li2CO3 production (Fig. 4a and Supplementary Fig. 17).

Figure 4b illustrates the mechanisms of the bipolar membrane in situ crystallization process, culminating in the formation of white powder products. Scanning electron microscopy (SEM) reveals that the precipitated powders exhibit a regular octahedral crystal structure with an average diameter of 4 μm (Fig. 4c and Supplementary Fig. 18). The element mapping confirms the homogeneous distribution of C and O within the particles (Supplementary Fig. 18). X-ray photoelectron spectroscopy (XPS) further vindicates the crystal purity and shows atomic percentages of 29.9% Li and 45.1% O, with less than 0.1% Na (Fig. 4d). High-resolution XPS spectra of C1s and O1s confirm the presence of C=O and C–O bonds in those generated crystals (Supplementary Fig. 19). X-ray diffraction (XRD) analysis further confirms that all the crystals are Li2CO3, demonstrating a high product purity (Fig. 4e).

The temporal evolution of Li+ and Na+ concentrations in the bipolar membrane in situ crystallization configuration is shown in Fig. 4f. At an optimized current density of 5 mA cm−2, both Li+ and Na+ in the feed chamber decrease to an extremely low value. In the recovery chamber, the Li+ initially increases and then decreases to 0.1 g L−1 due to the in situ crystallization of Li2CO3 (reaction mechanisms are detailed in Supplementary Fig. 17). This process yields Li2CO3 crystals with >99.9% purity (Fig. 4g). Without in situ crystallization, >99.0% of both Li+ and Na+ would be recovered (Fig. 4g and Supplementary Fig. 20). With this optimized configuration, >98.0% of Li+ is recovered from the leaching solution (achieves a recovery of 95.5% for Li+ in the whole process), and produces Li2CO3 with >99.9% purity. Additionally, the NaOH (>97.5% recovery, >95% purity) is recovered and reused as a basic solution in subsequent membrane metal-complexing ex situ crystallization as well as membrane metal-extracting temporal crystallization configurations (Supplementary Fig. 21). High-purity HCl is generated for spent LIB dissolution and subsequent systems for divalent cation recovery.

This bipolar membrane in situ crystallization system not only generates high-purity Li2CO3 but also provides significant energy, economic, and environmental benefits. The recovered Li2CO3, NaOH, and HCl mitigate environmental pollution of discharged metal ions and safety risks associated with high metal ion concentrations while offering valuable ion resources for industrial applications. Measuring by economic efficiency of per $ kg−1 Li2CO3, as shown in Fig. 4h, the economic analysis reveals that the primary costs are energy consumption (1.10 $ kg−1 Li2CO3) and investment (membranes: 1.68 $ kg−1 Li2CO3, maintenance: 0.43 $ kg−1 Li2CO3). Given the current market price of Li2CO3 (24.00 $ kg−1), the system yields a profit of around 20.79 $ kg−1 Li2CO3 (Supplementary Tables 10 and 11). Beyond the Li2CO3 profits, this configuration also generates NaOH and HCl for supporting other LIB recycling configurations, along with environmental benefits such as CO2 fixation and zero liquid discharge. Despite those advantages, the system faces limitations of the current electro-driven membranes due to the low ion throughput and poor anti-scaling properties (Supplementary Fig. 22 and Fig. 23). Future work should focus on improving electro-membrane performance to enhance the metal ion throughput and mitigate electro-membrane issues that enable more efficient and sustainable LIB recycling.

Recovery of Mn2+, Ni2+, and Co2+
The tandem leaching streams from the selective membrane dual-stage distillation system contain Mn2+, Ni2+, and Co2+, which are extracted as valuable metals from the black mass obtained after precise dismantling and separation of spent LIBs. To recover those metal ions, the membrane metal-complexing ex situ crystallization, as well as membrane metal-extracting temporal crystallization configurations, are developed, as illustrated in Fig. 5a. These systems exploit differential coordination behavior of Mn2+, Ni2+, and Co2+ to ethylenediaminetetraacetic acid (EDTA) and the extractant Cyanex 272 to selectively recover Mn3O4, Ni(OH)2, and Co(OH)2, respectively (Fig. 5b). According to coordination capacity to the EDTA complexing agent, in the membrane metal-complexing ex situ crystallization configuration, Ni2+ and Co2+ exhibits significantly higher coordination capacity with EDTA than Mn2+, thus forming [EDTA-Ni]2− (KNi = 2.51 × 1018) and [EDTA-Co]2− (KCo = 2.82 × 1016) (Supplementary Figs. 24 and 25). In contrast, Mn2+ exhibits a significantly lower coordination capacity with EDTA (KMn = 7.76 × 1013). Consequently, Mn2+ migrates to the recovery chamber as the experiment progresses (Supplementary Fig. 26f). A small amount of Ni2+ and Co2+ is also recovered along with Ni2+ when the molar ratio of EDTA to Ni+Co at 1, due to the incomplete complexation (Supplementary Figs. 26d, e). As the molar ratio of EDTA to Ni+Co increases to 1.05, no extra Ni2+ and Co2+ ions are observed in the recovery chamber, and Mn2+ approaches the highest recovery. However, a further increase in the molar ratio to 1.15 significantly decreases the recovery efficiency of Mn2+ due to the formation of Mn[EDTA]2−. Considering the higher recovery of Mn2+, the molar ratio of EDTA to Ni+Co is optimized at 1.05. With the extra-added H+, which is supplied from the bipolar membrane in situ crystallization, [EDTA-Ni]2− and [EDTA-Co]2− de-coordinate with EDTA, and the subsequent separation is carried out in the membrane metal-extracting temporal crystallization configuration. Based on the higher affinity capacity to Cyanex 272, the Co2+ is more likely to form complex Co(Cyanex 272)2 and remain in the solution (Supplementary Figs. 27 and 28). Meanwhile, the Ni2+ migrates to the recovery chamber under the electrical field and is finally recovered as the green crystal of Ni(OH)2 (Fig. 5cii). With temporal control, the pink Co(OH)2 is finally obtained from the extracted chamber of the membrane metal-extracting temporal crystallization system (Fig. 5ciii).

SEM imaging reveals distinct crystal morphologies with uniform elemental distribution for Mn, Ni, and Co, respectively (Fig. 5d). Various specific crystal structures with different diameters and high homogeneity of the elements’ distribution can be observed from those images. The XPS spectra results confirm high purity (>99.5%) of all obtained metal ion products (Fig. 5e), with high-resolution spectra validating bonding environments of Mn, Ni, and Co, respectively (Supplementary Figs. 29-31). The XRD analysis confirms phase-purity of the Mn3O4, Ni(OH)2, and Co(OH)2, respectively, with no other impurity peaks in each of the XRD spectra (Fig. 5f). The temporal evolution of Ni2+, Co2+. and Mn2+ concentrations in each chamber are measured using inductively coupled plasma optical emission spectrometry (ICP-OES). In the membrane metal-complexing ex situ crystallization configuration (Fig. 5g), the Ni2+ and Co2+ form [EDTA-Ni]2− and [EDTA-Co]2− and migrate to the Ni & Co collect chamber in the anode side, while Mn2+, which remains uncomplexed, transports toward the cathode into the Mn recovery chamber. Then [EDTA-Ni]2− and [EDTA-Co]2− are decomplexed using 2.0 M H2SO4, the collected Ni and Co solution is used in the next membrane metal-extracting temporal crystallization configuration (Fig. 5h). In the first 10 h, the Co2+ accumulates in the co-extraction chamber as the Co2+ complexes with Cyanex 272, forming non-charged Co(Cyanex 272)2, while positively charged Ni2+ migrates to the recovery chamber. After 10 h, Co(Cyanex 272)2 is dissociated into Co2+, which is transported to the recovery chamber later.

In the membrane metal-complexing ex situ crystallization system, with the optimized current density of 5 mA cm−2, the results reveal more than 99.5% Mn2+ recovery, and yield Mn3O4 with purity higher than 99.9% (Fig. 5i). Residual [EDTA-Ni]2− and [EDTA-Co]2− are retained in the collection chamber, ensuring fresh water reclamation and zero waste discharge. In the membrane metal-extracting temporal crystallization configuration, with optimized operating conditions, Ni2+ and Co2+ are recovered with recoveries of higher than 83.1% and 87.3%, generating the final products of Ni(OH)2 and Co(OH)2 with more than 99.5% and 92.5% purity, respectively (Fig. 5i).

The recovery of high-purity Mn3O4, Ni(OH)2, and Co(OH)2 not only addresses the heavy metal ion pollution but also delivers significant economic benefits. Taking Mn3O4 as an example, in the membrane metal-complexing ex situ crystallization configuration, taking the economic efficiency per $ kg−1 Mn3O4, the energy consumption and investment costs are 1.30 and 1.02 $ kg−1 Mn3O4, respectively (Fig. 5j). With a market price of ~9.00 $ kg−1 Mn3O4, this configuration yields profits of 6.68 $ kg−1 Mn3O4 (Supplementary Table 12). In the membrane metal-extracting temporal crystallization configuration, measured by economic efficiency of per $ kg−1 [Ni(OH)2 + Co(OH)2] (with the mole ratio of 47:53), the energy consumption and investment costs are 1.89 and 1.86 $ kg−1 [Ni(OH)2 + Co(OH)2], respectively (Fig. 5j). Given market prices of >14 $ kg−1 for Ni(OH)2 and >18 $ kg−1 for Co(OH)2, this system brings a profit at least 12.37 $ kg−1 [Ni(OH)2 + Co(OH)2] (Supplementary Table 13). Compared to the mainstream technologies for Li recovery, the e-MCGR technology not only achieves a lower overall cost but also offers distinct profit advantages (Supplementary Table 14). Besides these energy consumptions, the extra NaOH and HCl are supplied and reused from the bipolar membrane in situ crystallization system. The EDTA and the Cyanex 272 are recycled in membrane metal-complexing ex situ crystallization and membrane metal-extracting temporal crystallization systems, which minimize reagent consumption and enhance the system’s circularity. Considering the recovery rate of those metal ions, the profits could be improved when using the electro-driven membrane with high metal ion transport capacity in this configuration. Despite its energy, economic, and environmental advantages, those systems face challenges, particularly mineral scaling on electro-membranes, which limits metal ion throughput and recovery efficiency (Supplementary Fig. 32). To address these limitations, future efforts should focus on designing electro-membranes with higher ion transport capacity and enhanced anti-scaling properties. A fundamental study in advancing the understanding of metal ion crystallization kinetics in electro-membrane systems is essential to optimize recovery processes.

Recovery of metal ions from real LIBs leachate
The e-MCGR process for efficient LIBs recycling is further validated using real leachate from spent NMC111 cathode material (Fig. 6a). In the first-stage selective electrodialysis, operated at a current density of 1.0 mA cm−2, the CIMS membranes effectively reject multivalent metal ions (Mn2+, Ni2+, and Co2+) and permeate monovalent metal ions (Na+ and Li+). Thus, the concentrations of Li+ and Na+ increase to 1.08 g L−1 and 3.73 g L−1 in the recovery chamber, respectively (Fig. 6b). The residual solutions from the transitional chambers (in both transitional chamber-1 and transitional chamber-2) in the first-stage selective electrodialysis are then fed into the second-stage selective electrodialysis for further separation (Fig. 6c). This selective membrane dual-stage distillation configuration achieves accumulative recovery of 93.9% for Li+ and 89.1% for Na+, while completely rejecting Mn2+, Ni2+ and Co2+. This process demonstrates a low energy cost (1.56 $ kg−1), investment cost (7.08 $ kg−1), and energy consumption (14.87 kWh kg−1) (Supplementary Table 15).

Subsequent bipolar membrane in situ crystallization process recovers 96.6% Li+ from the leaching solution and produces Li2CO3 with > 99.9% purity (Fig. 6d, e). XPS analysis confirm the product’s composition, showing atomic percentages of 32.2% Li, and 39.6% O, with others less than 0.3% (Fig. 6f). The high purity of 99.9% was further verified by XRD analysis (Fig. 6g). The economic analysis reveals that the primary costs are energy consumption (0.72 $ kg−1 Li2CO3) and investment cost (2.12 $ kg−1 Li2CO3). Given the current market price of Li2CO3 (24.00 $ kg−1), the system yields a profit of around 21.16 $ kg−1 Li2CO3 (Supplementary Table 16). Moreover, 97.7% NaOH with 96.7% purity is recovered from the bipolar membrane in situ crystallization process.

In membrane metal-complexing ex situ crystallization, due to its lower coordination capacity with EDTA, Mn2+ migrates to the recovery chamber (Fig. 6h). In a separate membrane metal-extracting temporal crystallization step, Ni2+ migrates to the recovery chamber and is recovered as Ni(OH)2 by adding NaOH (Fig. 6i). With temporal control, the pink Co(OH)2 is subsequently obtained from the extraction chamber. The membrane metal-complexing ex situ crystallization proves optimal for maximizing the recovery rate and product purity for all transition metals (Fig. 6j). The XPS spectra results confirm high purity (>92.0%) of all obtained metal ion products (Fig. 6k). The XRD analysis confirms the generation of Mn3O4 (Fig. 6l). The energy consumption and investment costs are 1.42 and 1.46 $ kg−1 Mn3O4, respectively. With a market price of ~9.00 $ kg−1 Mn3O4, this configuration yields profits of 6.12 $ kg−1 Mn3O4 (Supplementary Table 17). XRD analysis confirmed the phase-purity of the recovered Ni(OH)2 and Co(OH)2 (Fig. 6m, n). The process achieved an 80.1% recovery of Ni(OH)2 with 99.3% purity and, with temporal control, a 95.6% recovery of pink Co(OH)2 with 92.2% purity. The combined energy consumption and investment costs for recovering Ni(OH)2 + Co(OH)2 are 1.61 and 1.93 $ kg−1 [Ni(OH)2 + Co(OH)2], respectively. Given market prices of >14 $ kg−1 for Ni(OH)2 and >18 $ kg−1 for Co(OH)2, this system brings a profit of 12.58 $ kg−1 [Ni(OH)2 + Co(OH)2] (Supplementary Table 18).

Methods
Pretreatment
The suspended solids in the feed are removed by using ultrafiltration (UP150 membrane) at an operating pressure of 0.01 MPa. Then the permeate is concentrated using reverse osmosis (RO, BW30 membrane) at an operating pressure of 1.5 MPa.

Configurations in the e-MCGR process
All experimental configurations described herein are conducted using a single, consistently prepared complex feed batch (the composition is shown in Supplementary Table 7) to ensure comparability and integration across the e-MCGR process. This batch simulated real LIB leachate is used sequentially across all experimental stages without intermediate reconstitution (Supplementary Table 19). The integrated design of the e-MCGR process comprises:

i. Selective membrane dual-stage distillation configuration: This configuration employs two sequential selective electro-membrane stacks to separate monovalent metal ions (Li+ and Na+) from the divalent metal ions (Mn2+, Ni2+, and Co2+). In the first-stage selective electro-membrane stack (Supplementary Fig. 6), the membranes are arranged as: anode / bipolar membrane (BM) / AMX (commercial anion exchange membrane) × 3 / CIMS (commercially monovalent cation exchange membrane) × 3 / bipolar membrane (BM) / cathode (Supplementary Table 20). The effective area per membrane is 0.0189 m2. Four chambers (feed, transitional-1, transitional-2, recovery) are connected to containers holding 2 L simulated LIB leaching solution in the feed chamber and 0.5 L deionized water in the other chambers. The operational parameters of the current densities are tested at 1 mA cm−2, 2.5 mA cm−2, and 5 mA cm−2. The electric current of the electro-membrane stack is supplied by a stabilized power supply. In order to further distill Li+ and Na+ from transitional chambers (transitional chamber-1 and transitional chamber-2) of the first stage, the second-stage selective electro-membrane stack is developed (Supplementary Fig. 6), and the electro-membrane stack arrangement is anode / BM / AMX × 2 / CIMS × 2 / BM / cathode. Three chambers in the electro-membrane stack consist of the feed chamber, the transitional chamber, and the recovery chamber. For those chambers, 1 L of residual solution from the first-stage transitional chambers is introduced into the feed chamber of the second-stage stacks, and the transitional and recovery chambers are initially fed with 0.5 L of deionized water. The current density of the electro-membrane stack is 0.5 mA cm−2. The refined Li+ and Na+ from the recovery chamber are then directly used as feed for the subsequent bipolar membrane in situ crystallization (ii) to separate Li+ from Na+.

ii. Bipolar membrane in situ crystallization configuration: In this configuration, the electro-membranes are stacked between anode and cathode in a sequence of the anode / BM / AMX / commercial cation exchange membrane (CMX) / BM / cathode (Supplementary Fig. 13), forming three chambers, including the feed chamber, the acid chamber, and the base chamber. Specifically, 1 L monovalent metal ions solution (Li+ and Na+) from the recovery chambers of configuration (i) is used as the initial feed, and 0.5 L of deionized water is fed to the acid and base chambers. The electric current of this electro-membrane stack is 5 mA cm−2, which is supplied by a direct current-stabilized power supply. To avoid Na+ impurities in Li2CO3 (caused by Na2CO3 co-precipitation) and capture the CO2 from industry, CO2 gas is introduced into the base chamber as a carbonate source, at a constant flow rate of 50 mL min−1 (Supplementary Fig. 14). During the reaction, CO2 dissolves into the solution and subsequently forms carbonate species. The Li2CO3 crystals will be crystallized in the base solution, followed by Li2CO3 recovery with the UP150 membrane.

CO2 + 2OH− ⇌ CO32− + H2O
Then, Li2CO3 crystallization occurs as
CO32− + 2Li+ ⇌ Li2CO3 (Ksp = 8.15 × 10−4)
Overall, the reaction is
2OH− + CO2 + 2Li+ ⇌ Li2CO3 ↓ + H2O

iii. Membrane metal-complexing ex situ crystallization configuration: Here, the membrane sequence is: anode / CMX / AMX / CMX / AMX / cathode (Supplementary Fig. 19), with three chambers: the feed, complexing, and the recovery chamber. The residual divalent metal-rich solution remaining after configuration (i) serves as the feed for this configuration. Herein, 3.0 L mixtures of Mn2+, Ni2+, and Co2+ are introduced into the feed chamber, and 0.5 L of deionized water is fed into the complexing and recovery chambers. The feed solution is further treated with the EDTA-4Na salt in the feed container. When the [EDTA]4− forms complex anions with Ni2+ and Co2+, with the EDTA’s differential coordination capacities (KNi > KCo ≫ KMn), the Mn2+ is separated from Ni2+ and Co2+ using this membrane metal-complexing ex situ crystallization system,

Ni2+ + EDTA4− ⇌ [EDTA-Ni]2− KNi = 2.51 × 1018
Co2+ + EDTA4− ⇌ [EDTA-Co]2− KCo = 2.82 × 1016
Mn2+ + EDTA4− ⇌ [EDTA-Mn]2− KMn = 7.76 × 1013

For the decomplexation, EDTA complexes are dissociated by adding 2.0 M HCl until pH < 0.5. To recover metal ions, Mn2+ in the recovery chamber is precipitated as Mn(OH)2 by using 1.0 M NaOH. Both the acid and base used here are provided by configuration (ii).

Mn2+ + 2OH− ⇌ Mn(OH)2 ↓
Then, the Mn(OH)2 is oxidized to black Mn3O4 crystals
6Mn(OH)2 + O2 ⇌ 2Mn3O4 ↓ + 6H2O

Membrane metal-extracting temporal crystallization configuration: Connecting with the membrane metal-complexing ex situ crystallization configuration, this stack uses: anode / AMX / CMX / CMX / AMX / cathode (Supplementary Fig. 19). In this system, the membrane stack contains three containers, which connect with the feed chamber, the Co-extraction chamber, and the Ni-recovery chamber. The containers are initially fed with 0.5 L mixed Ni2+ and Co2+ solution from the membrane metal-complexing ex situ crystallization system, 0.5 L extraction solution (50% saponified Cyanex 272 in kerosene), and 0.5 L deionized water, respectively. The electric current of the complexation electrodialysis stack is tested at 5 mA cm−2. Due to the Cyanex 272 preferentially binds Co2+, it enables efficient Ni2+ and Co2+ separation. In this process, the Co2+ is extracted by Cyanex 272, which forms a stable complex with Cyanex 272 in the extraction chamber.

Co2+ + 2Cyanex 272 ⇌ Co(Cyanex 272)2 + 2H+
Then, the uncomplexed Ni2+ migrates to the Ni-recovery chamber under the electric field. The Ni2+ is precipitated as Ni(OH)2 with the usage of 1.0 M NaOH (supplied from the bipolar membrane in situ crystallization system).

Ni2+ + 2OH− ⇌ Ni(OH)2 ↓
With temporal crystallization of Ni2+, 2.0 M HCl is added to the feed chamber to release Co2+ from Cyanex 272,
Co(Cyanex 272)2 + 2H+ ⇌ Co2+ + 2Cyanex 272
Then, these Co2+ migrate to the Co-recovery chamber under the electric field and precipitate as Co(OH)2 with the usage of 1.0 M NaOH.

Co2+ + 2OH− ⇌ Co(OH)2 ↓

Validation of the e-MCGR process using real LIBs wastewater
To validate the practical feasibility of the e-MCGR process, real wastewater from spent ternary LIBs is used as feed. The spent ternary LIBs, sourced from a technology company’s power bank, are used as the primary raw materials. First, 0.4 g of waste NMC111 cathode material is leached in 4.0 M HCl at 80 °C for 6 h. The resulting leachate is subjected to the same pretreatment sequence as described in the “Pretreatment” section, namely ultrafiltration (UP150 membrane, 0.01 MPa) to remove organic impurities, followed by RO (BW30 membrane, 1.5 MPa) to concentrate the solution. The concentrated leachate, now enriched in Li⁺, Na⁺, Mn²⁺, Ni²⁺, and Co²⁺, is directly used as the feed for the e-MCGR process, without reconstitution or modification, to ensure a one-to-one comparison with the results obtained from the simulated LIB leachate.

Characterization methods
The chemical composition and bonding states of the recovered crystals are analyzed by using X-ray photoelectron spectroscopy (XPS; AXIS Ultra DLD, Shimadzu, Japan). Scanning electron microscopy (SEM; Sirion 200, FEI, USA) coupled with energy-dispersive spectroscopy (EDS; INCA X-Act, USA) is employed to assess crystal morphology, surface roughness, and elemental distribution. The crystal phase identification is performed through X-ray diffractometry (XRD; XRD-6100, Shimadzu, Japan) using Cu-Kα radiation (λ=1.5 Å). Metal ion concentrations are quantified by using Inductively Coupled Plasma Optical Emission spectroscopy (ICP-OES; Avio 500, PerkinElmer, USA). The properties of ion exchange membranes used in the study are shown in Supplementary Table 12.

Evaluation methods
Technological analysis. Metal ion selectivity (P) is a parameter that is used to illustrate membrane ion selectivity. Compared to other metal ions (Mz+), the target metal ion selectivity can be calculated as
P = (Ttarget / Tother) / (Ctarget / Cother) * (Jtarget / Jother) (1)
Where Ttarget and Tother are the transport numbers of target metal ions and other metal ions. c (mol L−1) is the concentration of metal ions. J (mol m−2 s−1) is the metal ion permeation rate, which is used to illustrate ion transport capacity, and can be calculated as
J = (V * dc) / (S * dt) (2)
Where V (L) is the volume of solution, and S (m2) is the membrane area. The product recovery rate (R, %) is calculated by,
R = (mt / m) (3)
where mt (g) is the mass of precipitate recovered at time t, and m (g) is the theoretical maximum precipitate mass.
The equation of total suspended solids measurement, which is the crystal yield (Y, g L−1),
Y = mcry / Vbrine (4)
where mcry is the mass of dried crystals, and Vbrine (L) is the volume of feed wastewater.

Energy, economic, and environmental analysis. In this e-MCGR process, the specific energy consumption (E, kWh kg−1) is calculated by,
E = (∫0t Ut * I dt) / (Ct * Vt * Mt) (5)
where Ut (V) is the electro-membrane stack voltage, I (A) is the current, Vt (L) is the volume of solution, Ct (mol L−1) is the concentration, and Mt (g mol−1) is the molecular mass. More data on energy, economics, and the environment on this e-MCGR technology are analyzed and summarized in the Supplementary Information.

Data availability
The data generated in this study are provided in the article and its Supplementary Information. Visualization and figure preparation are carried out using Origin 8, Microsoft PowerPoint 2021, and Microsoft Excel 2021. Source data are provided with this paper.

Peer review rounds
[{"round": 1, "reviewer_comments": {"reviewer_1": "This study introduces a modular and innovative electro-membrane crystallization-assisted general recycling (e-MCGR) platform for spent lithium-ion batteries. The system integrates four key processes: dual-stage membrane separation of Li⁺/Na⁺ from multivalent cations, in-situ Li₂CO₃ crystallization with CO₂ utilization, ex-situ Mn₃O₄ recovery, and temporal crystallization for Ni(OH)₂ and Co(OH)₂. The configurations enable efficient recovery of high-purity metal products while minimizing chemical waste through the reuse of NaOH, HCl, and recycling of EDTA/Cyanex 272. The approach offers both environmental and economic benefits and is compatible with sustainable resource and energy systems. A major revision is, however, needed prior to considering for publication a prestigious journal like Nature Communications. The following are the comments that require the author's attention.\n1. Some of the presented results are unnecessary and repetitive. The authors should avoid repetitive schemes and adopt an effective data demonstration approach. For instance, Figure 1, Figure S1–S2, and Table S1–S4 provide the same information (i.e., data on lithium ion battery production and spent). Similarly, several unwanted schematics can also be avoided.\n2. Did the authors use waste Lithium batteries for the experiments, or were analytical grade chemicals utilized for the recovery demonstrations?\n3. The study demonstrated (Figure 2) the recovery of monovalent (Na+ and Li+) and divalent metals (Mn2+, Ni2+, Co2+). Did the authors use a single complex feed batch for all sets of experiments (i.e., recovery of Na+, Li+, Mn2+, Ni2+, and Co2+)? This is unclear, please specify.\n4. In Line 201 – 205: A cost of 1.24 $ Kg– 1 is reported for Li recovery at the first stage (RR = 87.7% at 1.0 mA). When adding the second stage, the cost rose to 2.73 $ Kg– 1. Was the cost of the second stage higher than the first stage, even though the used current density (0.5 mA cm – 2) was lower? Also, why did the authors use 0.5 mA? Was it optimized?\n5. In Lines 229–231, 100% rejection of multivalent cations is reported. Was this from the first stage, or is it the cumulative value for the two-stage process?\n6. The authors employed UF and RO for filtering macromolecules and concentrating the leaching solution for subsequent units. However, the details of the membrane types used and the operating conditions are not provided in the manuscript.\n7. Some of the terms used in the manuscript are speculative and misleading. The authors are advised to avoid using such terms.\n8. The provided methods are very descriptive. The authors are advised to include a detailed description of the methods to assist readership and support further investigations.\n9. The authors have reported the reuse of HCl, NaOH, and EDTA. Did the authors perform any treatment prior to reusing them? Up to how many repeated cycles was the performance unaffected?\n10. What price sources and methods did the authors use for performing the cost analysis?\n11. The manuscript requires thorough proofreading to correct typos and errors, such as those in Figure 4a, line 174, etc.", "reviewer_2": "I co-reviewed this manuscript with one of the reviewers who provided the listed reports. This is part of the Nature Communications initiative to facilitate training in peer review and to provide appropriate recognition for Early Career Researchers who co-review manuscripts.", "reviewer_3": "Zhao and his colleagues reported a combined electro-membrane crystallization process for spent lithium-ion batteries (LIBs) recycling, termed e-MCGR technology. The proposed technology incorporated four core modules: selective membrane dual-stage distillation, bipolar membrane in-situ crystallization, membrane metal-complexing ex-situ crystallization, and membrane metal-extracting temporal crystallization. However, similar individual module technologies have been proposed by peers for the recycling of lithium and other metallic elements from LIB streams (AIChE J. 68(2022), 17710; Nature Water 1(2023) 547–554; Chemical Engineering Journal 348(2018) 952–959; Resources, Conservation & Recycling 178 (2022) 106076; et al.). The novelty of this system is not clear because this article merely conceptually integrates these modular technologies. Furthermore, the above modules were studied separately in this manuscript, resulting in weak interconnectivity among them and insufficient comprehensive examination of the overall practical feasibility of the integrated technology. Additionally, the discussion on selective electrodialysis, electro-driven membrane in-sit crystallization and metal-extracting crystallization was insufficient, with some principal explanations containing scientific errors. The research results are insufficient to demonstrate the practical feasibility of the proposed e-MCGR technology. Therefore, this paper is not suitable for publication in Nature Communications. The detailed comments are as follows.\n1. The LIB leaching solution used in this article should be a simulated solution prepared based on the ion composition of an actual LIB leaching solution. The authors should use the actual waste LIB leaching solution to verify the effectiveness of the technology. In particular, the manuscript and the supplementary information did not provide water quality data for the LIB solution used, which is essential for evaluating the work carried out throughout the manuscript.\n2. The structure and content organization of the manuscript should be revised carefully to enhance its readability:\n(1) The discussion in lines 123-133 of the Results and Analysis section focuses on the importance of e-MCGR system development, which aligns more appropriately with the Introduction. Please justify this arrangement.\n(2) An in-depth theoretical analysis on the ion sieving mechanism of monovalent selective ion membranes (CIMS) is lacking.The authors repeatedly describe four ion sieving mechanisms (Lines 123-125, Lines161-164 and Figure S5) but fail to link these mechanisms to the experimental results. Since typically only 1–2 mechanisms are functionally relevant in conventional CIMS applications, specify the primary mechanism governing the CIMS in this study and its contribution to optimizing the design and operational efficiency of the dual-stage selective electrodialysis system.\n(3) Parameter equations in the manuscript were duplicated in those in the supplementary information, such as the product recovery rate, the total suspended solids measurement, the specific energy consumption.\n(4) The e-MCGR process configurations (Lines 418–503) overlap significantly with the membrane stack design and operating/testing descriptions in the supplementary information (Lines 111–193). Redundancy should be minimized.\n3. As depicted in FIGURE 4b, the Li2CO3 crystal formed in the base stream loop. However, the method and mechanism to prevent the deposition of Li2CO3 crystals within the membrane stack remain unexplained. Besides, for the in-situ crystallization module, the control strategy of the CO2 inlet flow rate should be clarified. It is recommended to provide a complete device photo.\n4. In the line 456 of the manuscript and the line 151 of the supplementary information, the Ksp of Li2CO3 should be about 0.0016 instead of 10^(-10.33).\n5. FIGURE S9 shows no permselectivity between Li+ and Ni2+ could not be observed, while the monovalent cation exchange membrane could reject most Ni2+ in Fig3 and FigS8, which is contradictory. The experimental results reflected in Figure S9 should be checked carefully, and the corresponding textual descriptions should be revised to aid readers in comprehending the ideas the authors intend to convey.\n6. In line 214-216, the phenomenon of mineral scaling on the membranes at high current density of 5 mA/cm2 was attributed to the poor anti-scaling properties of existing membranes. This explaination is unreasonable. As shown in FIGURE S12, the scaling phenomenon on the surface of the CIMS membrane was more severe than that of the AMX membrane. This should be attributed to the occurrence of water splitting on the surface of CIMS. Due to that most divalent ions are repelled by CIMS and are unable to serve as current carriers, water would be splitted as H+ and OH- to supplement the lack of current carrier flux at the cross section of CIMS, which would be more obvious at high current density.\n7. The electro-membrane processes operated at constant current densities. However, neither the manuscript nor the supplementary information reflected any voltage curves of membrane stacks during the operation. These curves are crucial for assessing process stability and concentration polarization. Relevant data and discussion must be added.\n8. The influence of the molar ratio of EDTA to Ni2+ and Co2+ was not discussed in the manuscript. However, the molar ratio is a key parameter determining the speciation of Ni2+, Co2+ and Mn2+ in the wastewater. A lower molar ratio is not conducive to the full conversion of Ni2+ and Co2+ into [EDTA-Ni]2- and [EDTA-Co]2- forms, which in turn is not conducive to the recovery of Ni2+ and Co2+, and will result in the recovered Mn2+ concentrate containing Ni2+ and Co2+ impurities. An excessively high molar ratio will turn part of Mn2+ into [EDTA-Mn]2- form, thereby reducing the recovery rate of Mn2+ and causing the recovered Ni2+ and Co2+ concentrates to contain Mn2+ impurities. Therefore, relevant results and discussion should be supplemented.\n9. In FIGURE 5h, the horizontal axes are labeled as \"Feed chamber\", \"Feed chamber\", and \"Recovery chamber\", respectively. However, according to FIGURE S25, they should be \"Feed chamber\", \"Co-extraction chamber\", and \"Recovery chamber\", respectively. Meanwhile, the concentration evolution of Ni2+ and Co2+ in Co-extraction chamber and Recovery chamber was confusing. Within 0 to 10 hours, the concentration of Co2+in the Co-extraction chamber continued to rise, while the concentration of Co2+in the Recovery chamber remained at a low level. This was mainly due to the complexion of Co2+ to Cyanex272 formed the non charged Co(Cyanex272)2. After 10 h, the Co2+ concentration in the Co-extraction chamber declined while that in recovery chamber increased with time. It seemed that H2SO4 was added into the feed solution to substitute Co2+ from Co(Cyanex272)2 (line 184-187 in supplementary information) at the operation time of 10h, but the manuscript did not explicitly clarify this point.\n10. FIGURE 5i indicates that the Ni2+ recovery rate at 10h was lower than 80%, but authors claimed that the recovery rate could reach 92.5%. Please provide a reasonable explanation.\n11. No feasible method for recycling EDTA and Cyanex 272 was presented or validated. Given their cost and environmental impact, reagent recycling is crucial for the feasibility of divalent ion separation methods.\n12. A puzzling phenomenon in economic analysis part is that, as a high-value chemical, Li2CO3 was less profitable than nickel and cobalt compounds in Table S7 and Table S9. Also, the cost of EDTA and Cyanex 272 was ignored in the economic analysis. Authors should clarify these issues and supplement the comparison with mainstream technologies for lithium recovery.\n13. In line 226-268 of the supplementary information, the methods of environmental analysis, economic analysis and energy analysis were described vaguely, without necessary data and evaluation methods to support these analysis.\n14. The FIGURE 1 in manuscript was not a core content and should not be included in the main text. Also, the cited literature was less relevant to the research content.\n15. In line 360-361, authors claimed that \"the extra added NaOH and HCl are supplied and reused from the bipolar membrane in-situ crystallization system\". It lacked necessary data support from a continuous integrated e-MCGR process."}, "author_response": "We sincerely thank the reviewer for the constructive evaluation of our work and for recognizing the environmental benefits and economic potential of the proposed electro-membrane crystallization-assisted general recycling (e-MCGR) platform. We appreciate the valuable feedback and agree that addressing the identified points will significantly strengthen the manuscript. We have carefully considered all comments and have substantially revised the manuscript.\n1. Some of the presented results are unnecessary and repetitive. The authors should avoid repetitive schemes and adopt an effective data demonstration approach. For instance, Figure 1, Figure S1-S2, and Table S1-S4 provide the same information (i.e., data on lithium ion battery production and spent). Similarly, several unwanted schematics can also be avoided.\nResponse:\nThank you very much for your suggestions. We have carefully reviewed the manuscript and supplementary information to minimize redundancy and ensure a more effective presentation of the data.\n(1) While we acknowledge the concern regarding overlapping content, we would like to clarify that Fig. 1 and Figs. S1-S2, as well as Table S1-S4, serve distinct purposes. Specifically, Fig. 1a and Fig. S1 depict global lithium-ion battery (LIB) production for different years (2024 and 2023, respectively). Likewise, Fig. 1b and Fig. S2 show the distribution of spent LIBs for the years 2024 and 2023, respectively. Our intention in including both years of data was to emphasise the rapid growth in LIB production and consumption, as well as the accumulation of spent LIBs, within a short time frame. This, we believe, strengthens the urgency and relevance of the study. In addition, Tables S1-S4 are included to indicate the data sources used in our analysis. These tables provide transparency and traceability, which we believe are essential for scientific rigour and reproducibility.\n(2) We deleted the redundant equations from the Supplementary Information, including those for product recovery rate, total suspended solids measurement, and specific energy consumption. We have also removed the membrane stack design and operating/testing descriptions from the Supplementary Information to avoid overlap with the e-MCGR process configurations presented in the main manuscript.\n2. Did the authors use waste Lithium batteries for the experiments, or were analytical grade chemicals utilized for the recovery demonstrations?\nResponse:\nThank you for your comment. Yes, in our initial experiments, we used a simulated leaching solution that mimics the composition of real spent lithium-ion battery leachates, prepared from analytical-grade chemicals. The composition of the simulated leaching solution has been added in Table S7. This approach allowed us to systematically evaluate and optimize the electro-membrane crystallization-assisted general recycling (e-MCGR) platform under controlled conditions.\nTable S7. Composition of simulated spent LIBs wastewater.\nComponents UF permeate (mg L-1) RO concentrate (mg L-1)\nLi+ 70 297\nNa+ 230 1104\nNi2+ 194 973\nCo2+ 195 972\nMn2+ 182 903\nIn addition, in response to the reviewer’s valuable suggestion on “using waste lithium-ion batteries for the experiments”, we have now conducted additional experiments using the real leaching solution derived from spent lithium-ion batteries (see validation of the e-MCGR process using actual LIBs wastewater in methods). The results, now included in the revised manuscript (Fig. 6 and Table S15), show that the Li+ recovery rate approaches 93.6% with product purity > 99.9%. This further validates the effectiveness and applicability of the e-MCGR platform for practical LIB recycling scenarios.\nTo validate the practical feasibility of the e-MCGR process, real wastewater from spent ternary LIBs is used as feed (Table S15). The spent ternary LIBs, sourced from a technology company's power bank, are used as the primary raw materials. First, 0.4 g of waste NMC111 cathode material is leached in 4.0 M HCl at 80 ºC for 6 h. The resulting leachate is treated by ultrafiltration (UP150 membrane, 0.01 MPa) to remove organic impurities, followed by RO (BW30 membrane, 1.5 MPa) to concentrate the solution. The e-MCGR process for efficient LIBs recycling is further validated using real leachate from spent NMC111 cathode material (Fig. 6a). In the first-stage selective electrodialysis, operated at a current density of 1.0 mA cm−2, the CIMS membranes effectively reject multivalent metal ions (Mn2+, Ni2+, and Co2+) and permeate monovalent metal ions (Na+ and Li+). Thus, the concentrations of Li+ and Na+ increase to 1.08 g L-1 and 3.73 g L-1 in the recovery chamber, respectively (Fig. 6b). The residual solutions from the transitional chambers (in both transitional chamber-1 and transitional chamber-2) in the first-stage selective electrodialysis are then fed into the second-stage selective electrodialysis for further separation (Fig. 6c). This selective membrane dual-stage distillation configuration achieves accumulative recovery of 93.9% for Li+ and 89.1% for Na+, while completely rejecting Mn2+, Ni2+ and Co2+. This process demonstrated the low energy cost (1.56 $ kg-1), investment cost (7.08 $ kg-1), and energy consumption (14.87 kWh kg-1) (Table S15).\nSubsequent bipolar membrane in-situ crystallization process recovers 96.6% Li+ from the leaching solution and produces Li CO with > 99.9% purity (Figs. 6d-e). XPS analysis confirm the product's composition, showing atomic percentages of 32.2% Li, and 39.6% O, with others less than 0.3% (Fig. 6f). The high-purity of 99.9% was further verified by XRD analysis (Fig. 6g). The economic analysis reveals that the primary costs are energy consumption (0.72 $ kg-1 Li CO ) and investment cost (2.12 $ kg-1 Li CO ). Given the current market price of Li CO (24.00 $ kg-1), the system yields a profit of around 21.16 $ kg-1 Li CO (Table S16). Moreover, 97.7% NaOH with 96.7% purity is recovered from the bipolar membrane in-situ crystallization process.\nIn membrane metal-complexing ex-situ crystallization, due to its lower coordination capacity with EDTA, Mn2+ migrates to the recovery chamber (Fig. 6h). In a separate membrane metal-extracting temporal crystallization step, Ni2+ migrates to the recovery chamber and is recovered as Ni(OH) by adding NaOH (Fig. 6i). With temporal control, the pink Co(OH) is subsequently obtained from the extraction chamber. The membrane metal-complexing ex-situ crystallization proves optimal for maximizing the recovery rate and product purity for all transition metals (Fig. 6j). The XPS spectra results confirm high purity (> 92.0%) of all obtained metal ion products (Fig. 6k). The XRD analysis confirms the generation of Mn O (Fig. 6l). The energy consumption and investment costs are 1.42 and 1.46 $ kg-1 Mn O , respectively. With a market price of ~ 9.00 $ kg-1 Mn O , this configuration yields profits of 6.12 $ kg-1 Mn O (Table S17). XRD analysis confirmed the phase purity of the recovered Ni(OH)₂ and Co(OH)₂ (Figs. 6m-n). The process achieved an 80.1% recovery of Ni(OH)₂ with 99.3% purity and, with temporal control, a 95.6% recovery of pink Co(OH)₂ with 92.2% purity. The combined energy consumption and investment costs for recovering Ni(OH)₂ + Co(OH)₂ are 1.61 and 1.93 $ kg-1 [Ni(OH) + Co(OH) ], respectively. Given market prices of > 14 $ kg-1 for Ni(OH) and > 18 $ kg-1 for Co(OH) , this system brings a profit of 12.58 $ kg-1 [Ni(OH) + Co(OH) ] (Table S18).\nFig. 6. Recovery of lithium and metal salts from real LIBs leaching solutions. a, Photo of spent LIBs leachate; b, First-stage selective electrodialysis separation of monovalent cations (Li+ and Na+) from divalent cations (Mn2+, Ni2+, and Co2+) under the current density of 1.0 mA cm−2; c, Second-stage selective electrodialysis separation of transitional solution for enhanced monovalent selectivity; d, Temporal evolution of Li+ and Na+ concentrations under current density of 5 mA cm-2 in bipolar membrane in-situ crystallization process; e, Recovery rates and product purity; f, XPS spectra showing atomic compositions of Li CO crystals (Li : 30.2%, O: 44.8%, C: 24.8%); g, XRD spectra of the synthesized crystals; h, Temporal evolution of Mn2+, Ni2+, and Co2+ concentrations in membrane metal-complexing ex-situ crystallization system; i, Temporal evolution of Ni2+ and Co2+ concentrations in membrane metal-extracting temporal crystallization system; j, Recovery and product purity of Mn O (> 94.5% recovery, >99.8% purity), Ni(OH) (> 80.1% recovery, > 99.3% purity), and Co(OH) (> 95.6% recovery, > 92.2% purity); k, XPS spectra showing atomic compositions of (i) Mn O (Mn: 27.4%, O: 31.3%), with < 0.3% impurities, (ii) Ni(OH) (Ni: 29.6%, O: 46.3%), and (iii) Co(OH) (Co: 24.8%, O: 42.4%), with < 1.0% impurities; l, XRD spectra confirming phase-pure Mn O (PDF#24-0734) powder; m, XRD spectra confirming phase-pure Ni(OH) (PDF#14-0117) powder; n, XRD spectra confirming phase-pure Co(OH) (PDF#30-0443) powder.\nTable S15. Composition of real spent LIBs wastewater.\nComponents UF permeate (mg L-1) RO concentrate (mg L-1)\nLi+ 67 298\nNa+ 229 1104\nNi2+ 171 853\nCo2+ 161 804\nMn2+ 132 653\nSupplementary Table S16. Performance analysis of dual-stage selective electrodialysis for real spent LIBs wastewater recovery.\nParameters Values\nEnergy consumption (EC) 14.87\nEnergy cost (C ) 1.56\ne\nInvestment cost (C) 7.08\ni\nLi+ recovery cost (C ) 8.64\nLi\nRecovery rate (R ) 93.9\nm\nSupplementary Table S17. Performance analysis of bipolar membrane in-situ crystallization for real spent LIBs wastewater recovery.\nParameters Values\nEnergy consumption (EC) 14.87\nEnergy cost (C ) 0.72\ne\nInvestment cost (C) 2.12\ni\nLi+ recovery cost (C ) 21.16\nLi\nRecovery rate (R ) 96.6\nm\nPurity (P ) 99.90\nm\nSupplementary Table S18. Performance analysis of membrane metal-complexing ex-situ crystallization for real spent LIBs wastewater recovery.\nParameters Values\nEnergy consumption (EC) 13.44\nEnergy cost (C ) 1.42\ne\nInvestment cost (C) 1.46\ni\nMn2+ recovery cost (C ) 2.86\nMn\nRecovery rate (R ) 94.7\nm\nPurity (P ) 99.8\nm\nSupplementary Table S19. Performance analysis of membrane metal-extracting temporal crystallization for real spent LIBs wastewater recovery.\nParameters Values\nEnergy consumption (EC) 15.30\nEnergy cost (C ) 1.61\ne\nInvestment cost (C) 1.93\ni\nNi2++Co2+ recovery cost (C ) 3.53\nLi\nNi(OH) recovery rate (R ) 80.10\n2 m, Ni\nCo(OH) recovery rate (R ) 95.61\n2 m, Co\nNi(OH) purity (P ) 99.30\n2 m, Ni\nCo(OH) purity (P ) 92.20\n2 m, Co\n3. The study demonstrated (Figure 2) the recovery of monovalent (Na+ and Li+) and divalent metals (Mn2+, Ni2+, Co2+). Did the authors use a single complex feed batch for all sets of experiments (i.e., recovery of Na+, Li+, Mn2+, Ni2+, and Co2+)? This is unclear, please specify.\nResponse:\nYes, during the experiment, we used a complex solution as a starting feed for the first stage separation of monovalent cations from divalent cations. Afterwards, the feed solution in the next stage always came from either the transitional chamber or the recovery chamber from the previous stage. We are sorry for the unclear description. We have revised the description of the e-MCGR process to make sure it is clear.\nWe have revised it in the manuscript in\nLines 496-500: All experimental configurations described herein were conducted using a single, consistently prepared complex feed batch (the composition is shown in Supplementary Table S7) to ensure comparability and integration across the e-MCGR process. This batch simulated real LIB leachate and was used sequentially across all experimental stages without intermediate reconstitution.\nLines 516-521: For those chambers, 1 L of residual solution from the first-stage transitional chambers is introduced into the feed chamber of the second-stage stacks, and the transitional and recovery chambers are initially fed with 0.5 L of deionized water. The current density of the electro-membrane stack is 0.5 mA cm-2. The refined Li+ and Na+ from the recovery chamber are then directly used as feed for the subsequent bipolar membrane in-situ crystallization (ii) to separate Li+ from Na+.\nLines 542-545: The residual divalent metal-rich solution remaining after configuration (i) serves as the feed for this configuration. Herein, 3.0 L mixtures of Mn2+, Ni2+, and Co2+ are introduced into the feed chamber, and 0.5 L of deionized water is fed into the complexing and recovery chambers.\n4. In Line 201 - 205: A cost of 1.24 $ Kg-1 is reported for Li+ recovery at the first stage (RR = 87.7% at 1.0 mA). When adding the second stage, the cost rose to 2.73 $ Kg– 1. Was the cost of the second stage higher than the first stage, even though the used current density (0.5 mA cm-2) was lower? Also, why did the authors use 0.5 mA? Was it optimized?\nResponse:\nThank you for your thoughtful questions.\n(1) Yes, the cost of the second-stage SED was higher than that of the first stage, despite the lower applied current density (0.5 mA cm⁻²). This is primarily due to two key factors: First, the ion concentration in the feed solution of the second-stage SED was much lower, which resulted in a higher ion migration resistance and thus a higher energy consumption required for ion transport. Second, the lower Li+ concentration in the second-stage feed resulted in a much smaller amount of recovered Li (8.2 kg Y-1) compared to 27.4 kg Y-1 in the first stage. In summary, the higher energy consumption and lower Li recovery amount in the second stage lead to a higher cost of $2.73 kg⁻¹ compared to $1.24 kg⁻¹ in the first stage.\n(2) Yes, the current density of 0.5 mA cm-2 was optimized. When selecting the current density in the second-stage SED, we tested the limiting current density of the CIMS membrane using the LIB wastewater, which was determined to be 5.7 mA cm-2. Thus, a low current density of 0.5 mA cm-2 can avoid concentration polarization throughout the process as the concentration of the feed continuously decreases during the experiment. A low current density of 0.5 mA cm-2 is beneficial for Li⁺ with Ni2+, Co2+, and Mn2+.\n5. In Lines 229-231, 100% rejection of multivalent cations is reported. Was this from the first stage, or is it the cumulative value for the two-stage process?\nResponse:\nThe 100% rejection of multivalent cations was the cumulative value of the dual-stage process. We also demonstrate in\nLines 174-178: “this selective membrane dual-stage distillation configuration proves optimal for separating Li+ and Na+ from Mn2+, Ni2+ and Co2+. With the combination of first-stage operating at 1 mA cm-2 and second-stage at 0.5 mA cm-2, this dual-stage distillation achieves the highest accumulative recoveries of 96.5% for Li+ and 96.4% for Na+, while completely rejecting Mn2+, Ni2+, and Co2+ (Fig. 3e and Fig. S8).”\nTo make it clear, we also indicate it in the caption of Fig. 3e (Lines 225-226): “the cumulative recovery rates of monovalent cations (Li+: 96.5%, Na+: 96.4%) and rejection rates of multivalent cations (Mn2+, Ni2+ and Co2+: 100%) from dual-stage distillation.”\n6. The authors employed UF and RO for filtering macromolecules and concentrating the leaching solution for subsequent units. However, the details of the membrane types used and the operating conditions are not provided in the manuscript.\nResponse:\nThank you very much for the comments. We have added the details of the membrane types used and the operating conditions in the Supporting Information.\nTable S5. Properties of UF membrane.\nMolecular Pure water\nWater contact\nMembrane Material weight cut-off permeability\nangle (º) c\n(kDa) a (LMH bar-1) b\nUP150 Polyethersulfone 150 501.7 70.1\na The data were provided by the manufacturers.\nb Measured with deionized water at 0.1 MPa.\nc Measured using a contact angle meter (Dropmeter A-200, MAIST Vision Inspection & Measurement Co., China).\nTable S6. Properties of RO membrane.\nPure water\nMolecular weight\nMembrane Material permeability (LMH\ncut-off (Da) a\nbar-1) b\nPolyamide thin film\nBW30 98 17.2\ncomposite\na The data were provided by the manufacturers.\nb Measured with deionized water at 0.5 MPa.\n7. Some of the terms used in the manuscript are speculative and misleading. The authors are advised to avoid using such terms.\nResponse:\nThank you very much for the comments. We have revised all the speculative and misleading terms throughout the manuscript. For example, (1) we revised the mechanism of monovalent selectivity of CIMS as only size exclusion and electrostatic repulsion (Fig. S5), instead of four mechanisms as previously discussed. “The polyamino compound layer grafted on the CIMS surface introduces abundant amino groups, imparting positive charges to the membrane surface. This positively charged layer strongly repels multivalent cations such as Ni²⁺, Co²⁺, and Mn²⁺, while monovalent cations (e.g., Na⁺, Li⁺) with lower charge density experience much weaker repulsion, enabling preferential transport. In addition, the relatively dense structure of the CIMS membrane matrix provides a size sieving effect, where Na⁺and Li⁺, with smaller hydrated radii, permeate more easily than the larger hydrated Ni2+, Co2+, and Mn2+. The synergy between electrostatic repulsion and size sieving enhances monovalent selectivity, thereby contributing to the optimized design and operational efficiency of the dual-stage selective electrodialysis system.” (2) We also completed the membrane scaling explanation. The severe scaling of the membrane can be explained by not only the poor anti-scaling properties of commercial electro-membranes themselves, but also the water splitting phenomenon (Fig. S13). More can be found in the manuscript with marks in red.\n8. The provided methods are very descriptive. The authors are advised to include a detailed description of the methods to assist readership and support further investigations.\nResponse:\nThanks very much for your advice. We have included more details in the methods, both in the main manuscript and supporting information, including configurations in the e-MCGR process for both simulated wastewater and actual wastewater, and cost analysis to help further investigation.\nMethods in the main manuscript:\nPretreatment\nThe suspended solids in the feed are removed by using ultrafiltration (UP150 membrane) at an operating pressure of 0.01 MPa. Then the permeate is concentrated using reverse osmosis (RO, BW30 membrane) at an operating pressure of 1.5 MPa.\nConfigurations in the e-MCGR process\nAll experimental configurations described herein were conducted using a single, consistently prepared complex feed batch (the composition is shown in Supplementary Table S7) to ensure comparability and integration across the e-MCGR process. This batch simulated real LIB leachate and was used sequentially across all experimental stages without intermediate reconstitution. The integrated design of the e-MCGR process comprises:\ni. Selective membrane dual-stage distillation configuration: This configuration employs two sequential selective electro-membrane stacks to separate monovalent metal ions (Li+ and Na+) from the divalent metal ions (Mn2+, Ni2+ and Co2+). In the first-stage selective electro-membrane stack (Fig. S6), the membranes are arranged as: anode / bipolar membrane (BM) / AMX (commercial anion exchange membrane) × 3 / CIMS (commercially monovalent cation exchange membrane) × 3 / bipolar membrane (BM) / cathode. The effective area per membrane is 0.0189 m2. Four chambers (feed, transitional-1, transitional-2, recovery) are connected to containers holding 2 L simulated LIB leaching solution in the feed chamber and 0.5 L deionized water in the other chambers. The operational parameters of the current densities are tested at 1 mA cm-2, 2.5 mA cm-2, and 5 mA cm-2. The electric current of the electro-membrane stack is supplied by a stabilized power supply. In order to further distill Li+ and Na+ from transitional chambers (transitional chamber-1 and transitional chamber-2) of the first stage, the second-stage selective electro-membrane stack is developed (Fig. S6), and the electro-membrane stack arrangement is anode / BM / AMX × 2 / CIMS × 2 / BM / cathode. Three chambers in the electro-membrane stack consist of the feed chamber, the transitional chamber, and the recovery chamber. For those chambers, 1 L of residual solution from the first-stage transitional chambers is introduced into the feed chamber of the second-stage stacks, and the transitional and recovery chambers are initially fed with 0.5 L of deionized water. The current density of the electro-membrane stack is 0.5 mA cm-2. The refined Li+ and Na+ from the recovery chamber are then directly used as feed for the subsequent bipolar membrane in-situ crystallization (ii) to separate Li+ from Na+.\nii. Bipolar membrane in-situ crystallization configuration: In this configuration, the electro-membranes are stacked between anode and cathode in a sequence of the anode / BM / AMX / commercial cation exchange membrane (CMX) / BM / cathode (Fig. S13), forming three chambers, including the feed chamber, the acid chamber, and the base chamber. Specifically, 1 L monovalent metal ions solution (Li+ and Na+) from the recovery chambers of configuration (i) is used as the initial feed, and 0.5 L of deionized water is fed to the acid and base chambers. The electric current of this electro-membrane stack is 5 mA cm-2, which is supplied by a direct current-stabilized power supply. To avoid Na+ impurities in Li CO (caused by Na CO co-precipitation) and capture the CO from industry, CO gas is introduced into the base chamber as a carbonate source, at a constant flow rate of 50 mL min-1 (Fig. S14). During the reaction, CO₂ dissolves into the solution and subsequently forms carbonate species. The Li CO crystals will be crystallized in the base solution, followed by Li CO recovery with the UP150 membrane.\niii. Membrane metal-complexing ex-situ crystallization configuration: Here, the membrane sequence is: anode / CMX / AMX / CMX / AMX / cathode (Fig. S19), with three chambers: the feed, complexing, and the recovery chamber. The residual divalent metal-rich solution remaining after configuration (i) serves as the feed for this configuration. Herein, 3.0 L mixtures of Mn2+, Ni2+, and Co2+ are introduced into the feed chamber, and 0.5 L of deionized water is fed into the complexing and recovery chambers.\nFor the decomplexation, EDTA complexes are dissociated by adding 2.0 M H SO until pH < 0.5. To recover metal ions, Mn2+ in the recovery chamber is precipitated as Mn(OH) by using 2.0 M NaOH. Both the acid and base used here are provided by configuration (ii).\nValidation of the e-MCGR process using actual LIBs wastewater\nTo validate the practical feasibility of the e-MCGR process, real wastewater from spent ternary LIBs is used as feed. The spent ternary LIBs, sourced from a technology company's power bank, are used as the primary raw materials. First, 0.4 g of waste NMC111 cathode material is leached in 4.0 M HCl at 80 ºC for 6 h. The resulting leachate is subjected to the same pretreatment sequence as described in Section “Pretreatment,” namely ultrafiltration (UP150 membrane, 0.01 MPa) to remove organic impurities, followed by RO (BW30 membrane, 1.5 MPa) to concentrate the solution. The concentrated leachate, now enriched in Li⁺, Na⁺, Mn²⁺, Ni²⁺, and Co²⁺, is directly used as the feed for the e-MCGR process, without reconstitution or modification, to ensure a one-to-one comparison with the results obtained from the simulated LIB leachate.\nSupplymary materials:\n1.2. Influence of the molar ratio of EDTA to Ni2+ and Co2+\nA range of molar ratios of EDTA-4Na to Ni2++Co2+ (i.e., 0.95, 1.0, 1.05, 1.1, and 1.15) is tested to maximize Mn2+ recovery efficiency in the recovery chamber. The current density of the complexation electrodialysis stack is 5 mA cm-2.\nMethod for recycling EDTA and Cyanex 272. In the e-MCGR system, HCl and NaOH can be recovered in the acid chamber and base chamber, respectively. The HCl concentration and NaOH concentration approached 1.03 mol/L and 1.1 mol/L, respectively. After the membrane metal-complexing ex-situ crystallization, Ni and Co were decomplexed from EDTA at pH below 1.5 by adding HCl. The EDTA solid was precipitated out and recovered according to the following reaction:\n[𝐸𝐷𝑇𝐴−𝑁𝑖]2−+2𝐻+→ 𝑁𝑖2++2𝐸𝐷𝑇𝐴 ↓\n[𝐸𝐷𝑇𝐴−𝐶𝑜]2−+2𝐻+→ 𝐶𝑜2++2𝐸𝐷𝑇𝐴 ↓\nAfter the membrane metal-extracting temporal crystallization, the Cyanex 272 was regenerated by adding HCl at pH below 1 according to the following reaction:\n2𝐻++𝐶𝑜(𝐶𝑦𝑎𝑛𝑒𝑥272) →2𝐶𝑦𝑎𝑛𝑒𝑥272+𝐶𝑜2+\n2\n2𝐻++𝑁𝑖(𝐶𝑦𝑎𝑛𝑒𝑥272) →2𝐶𝑦𝑎𝑛𝑒𝑥272+𝑁𝑖2+\n9. The authors have reported the reuse of HCl, NaOH, and EDTA. Did the authors perform any treatment prior to reusing them? Up to how many repeated cycles was the performance unaffected?\nResponse:\nThank you very much for your review. In our experiments, the HCl, NaOH, and EDTA solutions were directly reused without additional purification between cycles. The recovered HCl and NaOH have concentrations of 1.03 mol L-1 and 1.1 mol L-1, respectively. In membrane metal-complexing ex-situ crystallization, the EDTA can be precipitated out from [EDTA-Ni]2- and [EDTA-Co]2- in the complexing chamber by adding HCl.\nIn the membrane metal-extracting temporal crystallization, 51 mL of NaOH with a concentration of 1.10 mol L-1 is used for the precipitation of Co(OH) , and 55 mL of NaOH with a concentration of 1.10 mol L-1 is used for the precipitation of Ni(OH) . For the current experiment, this approach could sustain stable system performance for at least one repeated cycle. The method for recycling EDTA and Cyanex 272 is:\nIn the e-MCGR system, HCl and NaOH can be recovered in the acid chamber and base chamber, respectively. The HCl concentration and NaOH concentration approached 1.03 mol/L and 1.1 mol/L, respectively. After the membrane metal-complexing ex-situ crystallization, Ni and Co were decomplexed from EDTA at pH below 1.5 by adding HCl. The EDTA solid was precipitated out and recovered according to the following reaction:\n[𝐸𝐷𝑇𝐴−𝑁𝑖]2−+2𝐻+→ 𝑁𝑖2++2𝐸𝐷𝑇𝐴 ↓\n[𝐸𝐷𝑇𝐴−𝐶𝑜]2−+2𝐻+→ 𝐶𝑜2++2𝐸𝐷𝑇𝐴 ↓\nAfter the membrane metal-extracting temporal crystallization, the Cyanex 272 was regenerated by adding HCl at pH below 1 according to the following reaction:\n2𝐻++𝐶𝑜(𝐶𝑦𝑎𝑛𝑒𝑥272) →2𝐶𝑦𝑎𝑛𝑒𝑥272+𝐶𝑜2+\n2\n2𝐻++𝑁𝑖(𝐶𝑦𝑎𝑛𝑒𝑥272) →2𝐶𝑦𝑎𝑛𝑒𝑥272+𝑁𝑖2+\n2\n10. What price sources and methods did the authors use for performing the cost analysis?\nResponse:\nThank you very much for your review. The price sources of Li CO , Ni(OH) , Co(OH) , and Mn O are obtained from Alibaba.com (https://www.alibaba.com), which provides standardized and verified chemical pricing information from multiple suppliers.\nAdditionally, we have included the cost analysis model in the Supplementary Material:\nSupplementary Table S10. The economic analysis method.\nCost types\nEnergy cost\nEnergy consumption of product recovery Collected from the energy consumption of the\n(kg kWh-1) membrane stack\nEffective running time of membrane module\nper year (365 days × 24 h/day × 0.8) ÷ time\nProcess capacity per year (kg Y-1) consumed by membrane module operation\n(h) × resource recovery by membrane stack\noperation (kg)\nElectricity price ($ kWh-1) 0.10\nEnergy cost for resource recovery ($ kg- Electricity price ($ kWh-1) ÷ Energy\n1) consumption of product recovery (kg kWh-1)\nEnergy cost for peripheral equipment Energy cost for resource recovery ($ kg-1) ÷\n($ kg-1) 20\nEnergy cost for resource recovery ($ kg-1) +\nTotal energy cost ($ kg-1)\nEnergy cost for peripheral equipment ($ kg-1)\nInvestment cost\nEntire device lifetime (Y)a 3\nArea of ion exchange membrane (m2) ×price\nMembrane cost ($)\nof ion exchange membrane ($ m-2).\nMembrane equipment cost ($) Membrane cost ($) × 1.5\nAuxiliary equipment cost ($) Membrane equipment cost ($) × 0.7\nMembrane cost ($) + Membrane equipment\nTotal investment cost ($)\ncost ($)\nTotal investment cost × 1.083-Amortization\nAmortization per year ($ Y-1) per year ($ Y-1) × 1.08 - Amortization per year\n($ Y-1) ×1.082\nMaintenance cost per year ($ Y-1) Total investment cost ($) ÷ 10\nAmortization per year ($ Y-1) + Maintenance\nFixed cost per year ($ Y-1)\ncost per year ($ Y-1)\nFixed cost per year ($ Y-1) ÷ Process capacity\nFixed cost for product recovery ($ kg-1)\nper year (kg Y-1)\nFixed cost for product recovery ($ kg-1) + Total\nOverall product cost ($ kg-1)\nenergy cost ($ kg-1)\nProduct market price ($ kg-1) Collected from Alibaba.com\nPrice of product ($ kg-1) - Overall product cost\nTotal product profit ($ kg-1)\n($ kg-1)\n11. The manuscript requires thorough proofreading to correct typos and errors, such as those in Figure 4a, line 174, etc.\nResponse:\nWe are sorry for the typos and errors in the manuscript. We have carefully read the manuscript and revised all the typos and linguistic mistakes. For example, we have corrected the spelling errors in Figures 4 and 5.\nFig. 4. Sustainable lithium carbonate (Li CO ) production from monovalent cation\n2 3\nsolution and waste CO . a, Schematic of Li+ recovery via the bipolar membrane in-situ\n2\ncrystallization configuration and the mechanisms of Li CO formation through utilization; b,\n2 3\nOperational principles of the bipolar membrane in-situ crystallization system, with an inset\nshowing the collected Li CO powder (white); c, SEM images of the Li CO crystals (regular\n2 3 2 3\noctahedral morphology, ~4 µm average diameter) and element mapping of C and O\ndistribution; d, XPS survey spectra of the Li CO crystals, showing atomic percentages: Li\n2 3\n(24.9%), O (45.1%), negligible Na (<0.1%), and C (24.9%, with testing background); e, XRD\nspectra of the synthesized crystals, confirming the crystallographic phase of Li CO . All\n2 3\ndiffraction peaks align with the standard reference pattern for lithium carbonate (PDF#22-\n1141), demonstrating high phase purity and crystallinity.; f, Temporal evolution of Li+ and Na+\nconcentrations in feed and recovery chambers under optimized current density (5 mA cm-2),\nLi+ depletion correlates with Li CO crystallization; g, Recovery rates (Li+: 98.0%; Na+: 97.5%)\n2 3\nand product purity (Li CO : > 99.9%; NaOH: 95%); h, The energy, economic and\n2 3\nenvironmental analysis of the bipolar membrane in-situ crystallization system: energy cost\n(1.10 $ kg-1 Li CO ), investment (2.11 $ kg-1 Li CO ), and profits (20.79 $ kg-1 Li CO ) relative\n2 3 2 3 2 3\nto market price of Li CO (24.00 $ kg-1).\n2 3\nFig. 5. Recycling of Mn2+, Ni2+, and Co2+ from the divalent cation solutions. a, Schematic\nof the subsequent membrane metal-complexing ex-situ crystallization as well as membrane\nmetal-extracting temporal crystallization configurations for sequential separation of Mn2+, Ni2+,\nand Co2+; b, Mechanisms of the EDTA complexion (selective coordination with Ni2+ and Co2+)\nand the Cyanex 272 extraction (preferential binding to Co2+), enabling targeted metal recovery;\nc, Photographs of recovered products, (i) black Mn O , (ii) green Ni(OH) , and (iii) pink\n3 4 2\nCo(OH) powders; d, SEM images and elemental mapping of Mn O , Ni(OH) and Co(OH)\n2 3 4 2 2\ncrystals, respectively, confirming uniform morphology and homogeneous distribution of Mn,\nNi, and Co; e, XPS spectra showing atomic compositions of (i) Mn O (Mn: 26.9%, O: 37.6%),\n3 4\n(ii) Ni(OH) (Ni: 29.1%, O: 56.3%), and (iii) Co(OH) (Co: 23.7%, O: 54.9%), with <0.1%\n2 2\nimpurities; f, XRD spectra confirming phase-pure Mn O (PDF#24-0734), Ni(OH) (PDF#14-\n3 4 2\n0117), and Co(OH) (PDF#30-0443) powders, with no impurity peaks; g, Temporal evolution\n2\nof Mn2+, Ni2+, and Co2+ concentrations in membrane metal-complexing ex-situ crystallization\nsystem, achieving > 99.5% Mn2+ at 5 mA cm-2 and at EDTA : (Ni + Co) of 1.05; h, Temporal\nevolution of Ni2+ and Co2+ concentrations in membrane metal-extracting temporal\ncrystallization system, recovering > 83.1% Ni2+ and > 87.3% Co2+; i, Recovery and product\npurity of Mn O (>99.5% recovery, >99.9% purity), Ni(OH) (>83.1% recovery, >99.5% purity),\n3 4 2\nand Co(OH) (>87.3% recovery, >92.5% purity). j, The energy, economic, and environmental\n2\nanalysis: energy consumption (Mn O : 1.30 $ kg-1 Mn O ; Ni(OH) and Co(OH) : 1.89 $ kg-1\n3 4 3 4 2 2\n[Ni(OH) + Co(OH) ]), investment costs (Mn O : 1.02 $ kg-1 Mn O ; Ni(OH) and Co(OH) : 1.86\n2 2 3 4 3 4 2 2\n$ kg-1 [Ni(OH) + Co(OH) ]), and net profits (Mn O : 6.68 $ kg-1 Mn O ; Ni(OH) and Co(OH) :\n2 2 3 4 3 4 2 2\n12.37 $ kg-1 [Ni(OH) + Co(OH) ]).\n2 2\nReviewer: 2\nI co-reviewed this manuscript with one of the reviewers who provided the listed reports.\nThis is part of the Nature Communications initiative to facilitate training in peer review\nand to provide appropriate recognition for Early Career Researchers who co-review\nmanuscripts.\nResponse:\nThank you very much for your review and kind recommendation. The comments\nmentioned have been carefully answered.\nReviewer: 3\nZhao and his colleagues reported a combined electro-membrane crystallization\nprocess for spent lithium-ion batteries (LIBs) recycling, termed e-MCGR technology.\nThe proposed technology incorporated four core modules: selective membrane dual-\nstage distillation, bipolar membrane in-situ crystallization, membrane metal-\ncomplexing ex-situ crystallization, and membrane metal-extracting temporal\ncrystallization. However, similar individual module technologies have been proposed\nby peers for the recycling of lithium and other metallic elements from LIB streams\n(AIChE J. 68(2022), 17710; Nature Water 1(2023) 547-554; Chemical Engineering\nJournal 348(2018) 952-959; Resources, Conservation & Recycling 178 (2022) 106076;\net al.). (1) The novelty of this system is not clear because this article merely\nconceptually integrates these modular technologies. (2) Furthermore, the above\nmodules were studied separately in this manuscript, resulting in weak interconnectivity\namong them and insufficient comprehensive examination of the overall practical\nfeasibility of the integrated technology. (3) Additionally, the discussion on selective\nelectrodialysis, electro-driven membrane in-situ crystallization and metal-extracting\ncrystallization was insufficient, with some principal explanations containing scientific\nerrors. (4) The research results are insufficient to demonstrate the practical feasibility\nof the proposed e-MCGR technology. Therefore, this paper is not suitable for\npublication in Nature Communications. The detailed comments are as follows.\nResponse:\nWe sincerely appreciate the reviewer’s thorough evaluation of our manuscript and the\nvaluable feedback provided. We understand and respect the reviewer’s concerns\nregarding the novelty and practical feasibility of the proposed e-MCGR technology for\nrecycling spent lithium-ion batteries. We would like to address the main points raised:\n(1) While we acknowledge that similar individual module technologies have been\nreported previously (as cited by the reviewer), our work goes beyond simply presenting\nthese as an independent process. The novelty and significance of our study lie in: 1)\nas we mentioned in the introduction, the e-MCGR process is purposely designed to\naddress both separation precision and practical applicability by sequentially\nintegrating membrane crystallization, metal extraction, and metal complexation\nmodules. This integration enables stepwise separation of Li⁺, Na⁺, Mn²⁺, Ni²⁺, and Co²⁺\nat high recovery rates and product purities, while minimizing chemical waste through\nthe internal production and reuse of HCl, NaOH, and EDTA—features not\ndemonstrated in earlier single-module studies. 2) We conducted a systematic cost\nassessment based on actual market prices. This allows us to evaluate economic\nviability alongside technical performance, providing a more complete assessment of\nindustrial applicability. 3) Moreover, we have conducted additional experiments using\nthe actual leaching solution derived from real spent LIBs. The results, now included in\nthe revised manuscript, further validate the effectiveness and applicability of the\nproposed e-MCGR process for practical LIB recycling scenarios, which have not been\naddressed and validated in previous studies.\n(2) The e-MCGR process is experimentally implemented as an integrated sequence\n(as shown in Fig. 2), where the output stream from one module serves directly as the\ninput for the next without intermediate reconstitution. In the manuscript, each module\nis described separately in detail. This way helps to clarify its individual function and\noperating conditions, and to gain a deep understanding of the separation mechanism\nin each configuration.\nTo further enhance interconnectivity among different modules, we added more details\nto the “Configurations in the e-MCGR process” section to emphasize sequential\noperation rather than stand-alone experiments. Additionally, we present the real\nwastewater validation results in a way that reflects the cumulative performance of the\nintegrated system rather than only module-by-module outcomes (Fig. 6).\n(3) We have expanded the mechanistic discussion of each process and corrected\nscientific inaccuracies.\nSeparation of monovalent cations from divalent cations.\nThe mechanisms relevant to monovalent selectivity of CIMS have been revised in lines\n152-154. Based on the characteristics of the commercial CIMS used in this study, the\nprimary ion sieving mechanisms are identified as electrostatic repulsion and size\nsieving: “Recent advancements have enabled the use of commercial monovalent ion-\nselective electro-membranes, which operate based on mechanisms such as size\nexclusion and electrostatic repulsion (Fig. 3a and Fig. S5)”.\nThe explanation of membrane scaling has been revised in lines 205-213; “While\nselective membrane dual-stage distillation technology achieves a high separation\nefficiency, current commercial membranes are limited by their metal ion throughput.\nIn addition, they suffer from severe membrane scaling (Fig. S11 and Fig. S12). On the\none hand, this scaling can be attributed to the inherent poor anti-scaling properties of\ncommercial electro-membrane. On the other hand, water splitting may occur on the\nCIMS surface as most divalent ions are electrostatically repelled and cannot serve as\ncurrent carriers. The resulting shortage of charge carriers induces forced dissociation\nof water into H⁺ and OH⁻, with OH⁻ reacting with accumulated metal ions to form\ninsoluble hydroxide precursors that deposit on the membrane surface (Fig. S13).”\nRecovery of Mn2+, Ni2+, and Co2+\nMore discussion has been added regarding the influence of EDTA: (Ni + Co) molar\nratio on ion separation. In lines 319-331: “The influence of the EDTA: (Ni + Co) molar\nratio, including 0.95, 1.0, 1.05, and 1.15, on metal ion separation is explored (Fig. S26).\nThe concentration of Ni²⁺, Co²⁺ and Mn²⁺ in the feed chamber presented a decreasing\ntrend (Figs. S26a-c). The Ni²⁺ and Co²⁺ form Ni[EDTA]2+/Co[EDTA]2+ complex anions,\nthus they migrate towards the anode into the chamber. On the contrary, the Mn2+ \nmigrated towards the cathode into the recovery chamber as the experiment\nprogressed (Fig. S26f). A small amount of Ni2+ and Co2+ is also recovered along with\nNi2+ when the molar ratio of EDTA to Ni+Co at 1, due to the incomplete complexation\n(Figs. S26d-e). As the molar ratio of EDTA to Ni+Co increases to 1.05, no extra Ni2+ \nand Co2+ ions are observed in the recovery chamber, and Mn2+ approaches the\nhighest recovery. However, a further increase in the molar ratio to 1.15 significantly\ndecreased the recovery efficiency of Mn2+ due to the formation of Mn[EDTA]2+.\nConsidering the higher recovery of Mn2+, the molar ratio of EDTA to Ni+Co is optimized\nat 1.05.”\nMechanistic discussion of metal ion evolution in the membrane metal-extracting\ntemporal crystallization system has been added in\nLines 346-357: The temporal evolution of Ni2+, Co2+ and Mn2+ concentrations in each\nchamber was measured using inductively coupled plasma optical emission\nspectrometry (ICP-OES). In the membrane metal-complexing ex-situ crystallization\nconfiguration (Fig. 5g), the Ni2+ and Co2+ form [EDTA-Ni]2- and [EDTA-Co]2- and\nmigrate to the Ni & Co collect chamber in the anode side, while Mn2+, which remains\nuncomplexed, transports towards the cathode into the Mn recovery chamber. Then\n[EDTA-Ni]2- and [EDTA-Co]2- are decomplexed using 2.0 M H SO , the collected Ni\n2 4\nand Co solution is used in the next membrane metal-extracting temporal crystallization\nconfiguration (Fig. 5h). In the first 10 h, the Co2+ accumulates in the co-extraction\nchamber as the Co2+ complexes with Cyanex272, forming non-charged\nCo(Cyanex272) , while positively charged Ni2+ migrates to the recovery chamber.\n2\nchamber later.\n(4) To strengthen the demonstration of practicality, we have now included additional\nexperiments using the actual leaching solution derived from real spent LIBs (See\nValidation of the e-MCGR process using actual LIBs wastewater in methods). The\nresults, added in the revised manuscript (Fig. 6), confirm that the e-MCGR modules\ncan effectively recover key metal ions from actual spent LIBs leaching solution, with\nhigh recovery rates and high product purity for Li+, Ni2+, Mn2+, and Co2+. This further\nvalidates the effectiveness and applicability of the proposed e-MCGR process for\npractical LIB recycling scenarios. Besides, we provided a systematic cost techno-\neconomic analysis to show that this process not only enables zero waste discharge,\nbut also valorizes the ionic resources within spent LIB leachate through stepwise\nseparation and crystallization. The results highlight the process’s technical and\neconomic viability.\n1. The LIB leaching solution used in this article should be a simulated solution\nprepared based on the ion composition of an actual LIB leaching solution. The authors\nshould use the actual waste LIB leaching solution to verify the effectiveness of the\ntechnology. In particular, the manuscript and the supplementary information did not\nprovide water quality data for the LIB solution used, which is essential for evaluating\nthe work carried out throughout the manuscript.\nResponse:\nThanks for your comments. The detailed water quality parameters of our simulated\nLIB leaching solution have been provided in Table S7 of the Supplementary\nInformation. Moreover, we have conducted comprehensive validation experiments\nusing real spent LIB leaching solutions (Table S15).\nTable S7. Composition of simulated spent LIBs wastewater.\nSimulated solution UF permeate (mg L-1) RO concentrate (mg L-1)\nLi+ 70 297\nNa+ 230 1104\nNi2+ 194 973\nCo2+ 195 972\nMn2+ 182 903\nTable S15. Composition of real spent LIBs wastewater.\nReal LIBs wastewater UF permeate (mg L-1) RO concentrate (mg L-1)\nLi+ 67 298\nNa+ 229 1104\nNi2+ 171 853\nCo2+ 161 804\nMn2+ 132 653\nIn response to the valuable suggestion on “The authors should use the actual waste\nLIB leaching solution to verify the effectiveness of the technology”, we have now\nconducted additional experiments using the actual leaching solution derived from\nspent lithium-ion batteries (see validation of the e-MCGR process using actual LIBs\nwastewater in methods). The results, now included in the revised manuscript (Fig. 6\nand Table S15), show that the Li+ recovery rate approaches 93.6% with product purity >\n99.9%. This further validates the effectiveness and applicability of the e-MCGR\nplatform for practical LIB recycling scenarios.\n“To validate the practical feasibility of the e-MCGR process, real wastewater from\nspent ternary LIBs is used as feed (Table S15). The spent ternary LIBs, sourced from\na technology company's power bank, are used as the primary raw materials. First, 0.4\ng of waste NMC111 cathode material is leached in 4.0 M HCl at 80 ºC for 6 h. The\nresulting leachate is treated by ultrafiltration (UP150 membrane, 0.01 MPa) to remove\norganic impurities, followed by RO (BW30 membrane, 1.5 MPa) to concentrate the\nsolution. The e-MCGR process for efficient LIBs recycling is further validated using real\nleachate from spent NMC111 cathode material (Fig. 6a). In the first-stage selective\nelectrodialysis, operated at a current density of 1.0 mA cm−2, the CIMS membranes effectively\nreject multivalent metal ions (Mn2+, Ni2+, and Co2+) and permeate monovalent metal ions (Na+\nand Li+). Thus, the concentrations of Li+ and Na+ increase to 1.08 g L-1 and 3.73 g L-1 in the\nrecovery chamber, respectively (Fig. 6b). The residual solutions from the transitional\nchambers (in both transitional chamber-1 and transitional chamber-2) in the first-stage\nselective electrodialysis are then fed into the second-stage selective electrodialysis for further\nseparation (Fig. 6c). This selective membrane dual-stage distillation configuration achieves\naccumulative recovery of 93.9% for Li+ and 89.1% for Na+, while completely rejecting Mn2+,\nNi2+ and Co2+. This process demonstrated the low energy cost (1.56 $ kg-1), investment cost\n(7.08 $ kg-1), and energy consumption (14.87 kWh kg-1) (Table S15).\nSubsequent bipolar membrane in-situ crystallization process recovers 96.6% Li+ from the\nleaching solution and produces Li CO with > 99.9% purity (Figs. 6d-e). XPS analysis confirm\n2 3\nthe product's composition, showing atomic percentages of 32.2% Li, and 39.6% O, with others\nless than 0.3% (Fig. 6f). The high-purity of 99.9% was further verified by XRD analysis (Fig.\n6g). The economic analysis reveals that the primary costs are energy consumption (0.72 $ kg-\n2 3 2 3 2 3\n(24.00 $ kg-1), the system yields a profit of around 21.16 $ kg-1 Li CO (Table S16). Moreover,\n2 3\n97.7% NaOH with 96.7% purity is recovered from the bipolar membrane in-situ crystallization\nprocess.\nIn membrane metal-complexing ex-situ crystallization, due to its lower coordination capacity\nwith EDTA, Mn2+ migrates to the recovery chamber (Fig. 6h). In a separate membrane metal-\nextracting temporal crystallization step, Ni2+ migrates to the recovery chamber and is\nrecovered as Ni(OH) by adding NaOH (Fig. 6i). With temporal control, the pink Co(OH) is\n2 2\nsubsequently obtained from the extraction chamber. The membrane metal-complexing ex-situ\ncrystallization proves optimal for maximizing the recovery rate and product purity for all\ntransition metals (Fig. 6j). The XPS spectra results confirm high purity (> 92.0%) of all obtained\nmetal ion products (Fig. 6k). The XRD analysis confirms the generation of Mn O (Fig. 6l). The\n3 4\nenergy consumption and investment costs are 1.42 and 1.46 $ kg-1 Mn O , respectively. With\n3 4\na market price of ~ 9.00 $ kg-1 Mn O , this configuration yields profits of 6.12 $ kg-1 Mn O\n3 4 3 4\n(Table S17). XRD analysis confirmed the phase purity of the recovered Ni(OH)₂ and Co(OH)₂\n(Figs. 6m-n). The process achieved an 80.1% recovery of Ni(OH)₂ with 99.3% purity and, with\ntemporal control, a 95.6% recovery of pink Co(OH)₂ with 92.2% purity. The combined energy\nconsumption and investment costs for recovering Ni(OH)₂ + Co(OH)₂ are 1.61 and 1.93 $ kg-\n1 [Ni(OH) + Co(OH) ], respectively. Given market prices of > 14 $ kg-1 for Ni(OH) and > 18\n2 2 2\n$ kg-1 for Co(OH) , this system brings a profit of 12.58 $ kg-1 [Ni(OH) + Co(OH) ] (Table S18).\n2 2 2\nFig. 6. Recovery of lithium and metal salts from real LIBs leaching solutions. a, Photo of\nspent LIBs leachate; b, First-stage selective electrodialysis separation of monovalent cations\n(Li+ and Na+) from divalent cations (Mn2+, Ni2+, and Co2+) under the current density of 1.0 mA\ncm−2; c, Second-stage selective electrodialysis separation of transitional solution for enhanced\nmonovalent selectivity; d, Temporal evolution of Li+ and Na+ concentrations under current\ndensity of 5 mA cm-2 in bipolar membrane in-situ crystallization process; e, Recovery rates\nand product purity; f, XPS spectra showing atomic compositions of Li CO crystals (Li : 30.2%,\n2 3\nO: 44.8%, C: 24.8%); g, XRD spectra of the synthesized crystals; h, Temporal evolution of\nMn2+, Ni2+, and Co2+ concentrations in membrane metal-complexing ex-situ crystallization\nsystem; i, Temporal evolution of Ni2+ and Co2+ concentrations in membrane metal-extracting\ntemporal crystallization system; j, Recovery and product purity of Mn O (> 94.5%\n3 4\nrecovery, >99.8% purity), Ni(OH) (> 80.1% recovery, > 99.3% purity), and Co(OH) (> 95.6%\n2 2\nrecovery, > 92.2% purity); k, XPS spectra showing atomic compositions of (i) Mn O (Mn:\n3 4\n27.4%, O: 31.3%), with < 0.3% impurities, (ii) Ni(OH) (Ni: 29.6%, O: 46.3%), and (iii) Co(OH)\n2 2\n(Co: 24.8%, O: 42.4%), with < 1.0% impurities; l, XRD spectra confirming phase-pure Mn O\n3 4\n(PDF#24-0734) powder; m, XRD spectra confirming phase-pure Ni(OH) (PDF#14-0117)\n2\npowder; n, XRD spectra confirming phase-pure Co(OH) (PDF#30-0443) powder.\nTo validate the practical feasibility of the e-MCGR process, real wastewater from spent ternary\nLIBs is used as feed. The spent ternary LIBs, sourced from a technology company's power\nbank, are used as the primary raw materials. First, 0.4 g of waste NMC111 cathode material\nis leached in 4.0 M HCl at 80 ºC for 6 h. The resulting leachate is subjected to the same\npretreatment sequence as described in Section “Pretreatment,” namely ultrafiltration (UP150\nmembrane, 0.01 MPa) to remove organic impurities, followed by RO (BW30 membrane, 1.5\nMPa) to concentrate the solution. The concentrated leachate, now enriched in Li⁺, Na⁺, Mn²⁺,\nNi²⁺, and Co²⁺, is directly used as the feed for the e-MCGR process, without reconstitution or\nmodification, to ensure a one-to-one comparison with the results obtained from the simulated\nLIB leachate.\nTable S15. Composition of real spent LIBs wastewater.\nComponents UF permeate (mg L-1) RO concentrate (mg L-1)\nLi+ 67 298\nNa+ 229 1104\nNi2+ 171 853\nCo2+ 161 804\nMn2+ 132 653\nSupplementary Table S16. Performance analysis of dual-stage selective\nelectrodialysis for real spent LIBs wastewater recovery.\nParameters Values\nEnergy consumption (EC) 14.87\nEnergy cost (C ) 1.56\ne\nInvestment cost (C) 7.08\ni\nLi+ recovery cost (C ) 8.64\nLi\nRecovery rate (R ) 93.9\nm\nSupplementary Table S17. Performance analysis of bipolar membrane in-situ\ncrystallization for real spent LIBs wastewater recovery.\nParameters Values\nEnergy consumption (EC) 14.87\nEnergy cost (C ) 0.72\ne\nInvestment cost (C) 2.12\ni\nLi+ recovery cost (C ) 21.16\nLi\nRecovery rate (R ) 96.6\nm\nPurity (P ) 99.90\nm\nSupplementary Table S18. Performance analysis of membrane metal-complexing\nex-situ crystallization for real spent LIBs wastewater recovery.\nParameters Values\nEnergy consumption (EC) 13.44\nEnergy cost (C ) 1.42\ne\nInvestment cost (C) 1.46\ni\nMn2+ recovery cost (C ) 2.86\nMn\nRecovery rate (R ) 94.7\nm\nPurity (P ) 99.8\nm\nSupplementary Table S19. Performance analysis of membrane metal-extracting\ntemporal crystallization for real spent LIBs wastewater recovery.\nParameters Values\nEnergy consumption (EC) 15.30\nEnergy cost (C ) 1.61\ne\nInvestment cost (C) 1.93\ni\nNi2++Co2+ recovery cost (C ) 3.53\nLi\nNi(OH) recovery rate (R ) 80.10\n2 m, Ni\nCo(OH) recovery rate (R ) 95.61\n2 m, Co\nNi(OH) purity (P ) 99.30\n2 m, Ni\nCo(OH) purity (P ) 92.20\n2 m, Co\n2. The structure and content organization of the manuscript should be revised carefully\nto enhance its readability:\nResponse:\nThanks for your comments. In response, we have conducted a comprehensive\nrevision of the manuscript to improve its clarity and logical flow. Specifically:\n(1) We have combined the first paragraph in the results and analysis section with the\nintroduction to better articulate the motivation for sustainable LIB recycling and\nlogically lead into the proposed e-MCGR concept.\nLines 95-101: Previously, membrane technologies have relied on principles such as\nthe importance of e-MCGR system development, which aligns more appropriately with\nthe Introduction. Please justify this arrangement.\nResponse:\nThanks for your suggestions. We have integrated this section with the introduction to\ncreate a smooth conceptual bridge from the background to the motivation for the e-\nMCGR design. We have adapted it in the manuscript in\nLines 95-130: “Previously, membrane technologies have relied on principles such as\nsize exclusion, electrostatic interactions, hydrophilic-hydrophobic interactions, and\nbonding affinity to achieve ion selectivity.31,40-42 Although substantial progress has\nbeen made toward developing membranes with precise single-ion selectivity, their\nutility in real-world resource recovery remains limited by material constraints and has\nlargely been tested only in isolated/single systems.43-45 Furthermore, their applicability\nin LIB recycling has not been convincingly demonstrated.42,46\nTo bridge this gap, we propose an e-MC-assisted general recycling (e-MCGR) concept,\nwhich integrates mature membranes into a closed-loop and modular framework\ntailored for spent LIB recycling. The e-MCGR concept is designed to address both\nseparation precision and practical applicability while maintaining high recovery and\nproduct purity. As illustrated in Fig. 2, the process begins with disassembly and acidic\ndissolution of spent LIBs, followed by pH adjustment (~7) to generate a representative\nleaching solution.⁴⁷ Valuable metals ions, such as Li+, Mn2+, Ni2+ and Co2+, are then\nselectively recovered through a three-part process: (i) separation of monovalent\ncations (Li+ and Na+) from divalent cations (Mn2+, Ni2+ and Co2+) through selective\nmembrane dual-stage distillation; (ii), separation of monovalent cations from each\nother (separation Li+ and Na+) and Li+ crystallization via bipolar membrane in-situ\ncrystallization and (iii) selective recovery of Mn2+, Ni2+ and Co2+ by membrane metal-\ncomplexing ex-situ crystallization and membrane metal-extracting temporal\ncrystallization. These processes enable the recovery of high-purity Li CO , Mn O ,\n2 3 3 4\nNi(OH) , and Co(OH) . Despite the strategic importance of metal ion recovery from\n2 2\nspent LIBs, no comprehensive study or design has yet explored a fully sustainable LIB\nrecycling through this e-MCGR concept. Furthermore, the potential of LIB recycling to\nenhance energy security, environmental sustainability, and economic viability remains\ninsufficiently explored. This e-MCGR presents an innovative framework and practical\nstrategy for optimizing metal ion recovery, thereby addressing the challenges\nencountered in sustainable resource recovery from LIBs. Given the global scale of LIB\ndevelopment, this e-MCGR concept and approach contribute a holistic, scalable, and\nsustainable solution to address global challenges in LIB waste management and\nresource recovery.”\n(2) We have also expanded the mechanistic discussion of each process to provide a\nclearer explanation of the underlying principles and highlight the interconnections\nbetween sections. Separation of monovalent cations from divalent cations.\nThe mechanisms relevant to monovalent selectivity of CIMS have been revised in lines\n152-154. Based on the characteristics of the commercial CIMS used in this study, the\nprimary ion sieving mechanisms are identified as electrostatic repulsion and size\nsieving: “Recent advancements have enabled the use of commercial monovalent ion-\nselective electro-membranes, which operate based on mechanisms such as size\nexclusion and electrostatic repulsion (Fig. 3a and Fig. S5).”\nThe explanation of membrane scaling has been revised in\nLines 205-213: “While selective membrane dual-stage distillation technology achieves\na high separation efficiency, current commercial membranes are limited by their metal\nion throughput. In addition, they suffer from severe membrane scaling (Fig. S11 and\nFig. S12). On the one hand, this scaling can be attributed to the inherent poor anti-\nscaling properties of commercial electro-membrane. On the other hand, water splitting\nmay occur on the CIMS surface as most divalent ions are electrostatically repelled and\ncannot serve as current carriers. The resulting shortage of charge carriers induces\nforced dissociation of water into H⁺ and OH⁻, with OH⁻ reacting with accumulated\nmetal ions to form insoluble hydroxide precursors that deposit on the membrane\nsurface (Fig. S13).”\nRecovery of Mn2+, Ni2+, and Co2+\nMore discussion has been added regarding the influence of EDTA: (Ni + Co) molar\nratio on ion separation. In\nLines 319-331: “The influence of the EDTA: (Ni + Co) molar ratio, including 0.95, 1.0,\n1.05, and 1.15, on metal ion separation is explored (Fig. S26). The concentration of\nNi²⁺, Co²⁺ and Mn²⁺ in the feed chamber presented a decreasing trend (Fig. S26a-c).\nThe Ni²⁺ and Co²⁺ form Ni[EDTA]2+/Co[EDTA]2+ complex anions, thus they migrate\ntowards the anode into the chamber. On the contrary, the Mn2+ migrated towards the\ncathode into the recovery chamber as the experiment progressed (Fig. S26f). A small\namount of Ni2+ and Co2+ is also recovered along with Ni2+ when the molar ratio of\nEDTA to Ni+Co at 1, due to the incomplete complexation (Fig. S26d-e). As the molar\nratio of EDTA to Ni+Co increases to 1.05, no extra Ni2+ and Co2+ ions are observed in\nthe recovery chamber, and Mn2+ approaches the highest recovery. However, a further\nincrease in the molar ratio to 1.15 significantly decreased the recovery efficiency of\nMn2+ due to the formation of Mn[EDTA]2+. Considering the higher recovery of Mn2+,\nthe molar ratio of EDTA to Ni+Co is optimized at 1.05.”\nMechanistic discussion of metal ion evolution in the membrane metal-extracting\ntemporal crystallization system has been added in\nLines 346-357: The temporal evolution of Ni2+, Co2+ and Mn2+ concentrations in each\nchamber was measured using inductively coupled plasma optical emission\nspectrometry (ICP-OES). In the membrane metal-complexing ex-situ crystallization\nconfiguration (Fig. 5g), the Ni2+ and Co2+ form [EDTA-Ni]2- and [EDTA-Co]2- and\nmigrate to the Ni & Co collect chamber in the anode side, while Mn2+, which remains\nuncomplexed, transports towards the cathode into the Mn recovery chamber. Then\n[EDTA-Ni]2- and [EDTA-Co]2- are decomplexed using 2.0 M H SO , the collected Ni\n2 4\nand Co solution is used in the next membrane metal-extracting temporal crystallization\nconfiguration (Fig. 5h). In the first 10 h, the Co2+ accumulates in the co-extraction\nchamber as the Co2+ complexes with Cyanex272, forming non-charged\nCo(Cyanex272) , while positively charged Ni2+ migrates to the recovery chamber.\n2\nchamber later.”\n(3) Methods: To further enhance interconnectivity among different modules, we added\nmore details to the “Configurations in the e-MCGR process” section to emphasize\nsequential operation rather than stand-alone experiments. Additionally, we present the\nreal wastewater validation results in a way that reflects the cumulative performance of\nthe integrated system rather than only module-by-module outcomes (Fig. 6).\n2.1. The discussion in lines 123-133 of the Results and Analysis section focuses on\nthe importance of e-MCGR system development, which aligns more appropriately with\nthe Introduction. Please justify this arrangement.\nResponse:\nThanks for your suggestions. We have integrated this section with the introduction to\ncreate a smooth conceptual bridge from the background to the motivation for the e-\nMCGR design. We have adapted it in the manuscript in\nLines 95-130: “Previously, membrane technologies have relied on principles such as\nsize exclusion, electrostatic interactions, hydrophilic-hydrophobic interactions, and\nbonding affinity to achieve ion selectivity.31,40-42 Although substantial progress has\nbeen made toward developing membranes with precise single-ion selectivity, their\nutility in real-world resource recovery remains limited by material constraints and has\nlargely been tested only in isolated/single systems.43-45 Furthermore, their applicability\nin LIB recycling has not been convincingly demonstrated.42,46\nTo bridge this gap, we propose an e-MC-assisted general recycling (e-MCGR) concept,\nwhich integrates mature membranes into a closed-loop and modular framework\ntailored for spent LIB recycling. The e-MCGR concept is designed to address both\nseparation precision and practical applicability while maintaining high recovery and\nproduct purity. As illustrated in Fig. 2, the process begins with disassembly and acidic\ndissolution of spent LIBs, followed by pH adjustment (~7) to generate a representative\nleaching solution.⁴⁷ Valuable metals ions, such as Li+, Mn2+, Ni2+ and Co2+, are then\nselectively recovered through a three-part process: (i) separation of monovalent\ncations (Li+ and Na+) from divalent cations (Mn2+, Ni2+ and Co2+) through selective\nmembrane dual-stage distillation; (ii), separation of monovalent cations from each\nother (separation Li+ and Na+) and Li+ crystallization via bipolar membrane in-situ\ncrystallization and (iii) selective recovery of Mn2+, Ni2+ and Co2+ by membrane metal-\ncomplexing ex-situ crystallization and membrane metal-extracting temporal\ncrystallization. These processes enable the recovery of high-purity Li CO , Mn O ,\n2 3 3 4\nNi(OH) , and Co(OH) . Despite the strategic importance of metal ion recovery from\n2 2\nspent LIBs, no comprehensive study or design has yet explored a fully sustainable LIB\nrecycling through this e-MCGR concept. Furthermore, the potential of LIB recycling to\nenhance energy security, environmental sustainability, and economic viability remains\ninsufficiently explored. This e-MCGR presents an innovative framework and practical\nstrategy for optimizing metal ion recovery, thereby addressing the challenges\nencountered in sustainable resource recovery from LIBs. Given the global scale of LIB\ndevelopment, this e-MCGR concept and approach contribute a holistic, scalable, and\nsustainable solution to address global challenges in LIB waste management and\nresource recovery.”\n2.2. An in-depth theoretical analysis on the ion sieving mechanism of monovalent\nselective ion membranes (CIMS) is lacking. The authors repeatedly describe four ion\nsieving mechanisms (Lines 123-125, Lines161-164 and Figure S5) but fail to link these\nmechanisms to the experimental results. Since typically only 1–2 mechanisms are\nfunctionally relevant in conventional CIMS applications, specify the primary\nmechanism governing the CIMS in this study and its contribution to optimizing the\ndesign and operational efficiency of the dual-stage selective electrodialysis system.\nResponse:\nThanks for your valuable comment. We have revised the manuscript to make it clear.\n“Previously, membrane technologies have relied on principles such as size exclusion,\nelectrostatic interactions, and chemical affinity to achieve ion selectivity.31,40-42\nAlthough substantial progress has been made toward developing membranes with\nprecise single-ion selectivity, their utility in real-world resource recovery remains\nlimited by material constraints and has largely been tested only in isolated/single\nsystems.43-45 Furthermore, their applicability in LIB recycling has not been convincingly\ndemonstrated.”\nThe mechanisms for preparing electro-driven membrane with ion selectivity mainly\ninclude membrane pore size sieving, electrostatic repulsion, and chemical affinity.\nBased on the characteristics of the commercial CIMS used in this study:\nThe primary ion sieving mechanisms of CIMS are identified as electrostatic repulsion\nand size sieving. The polyamino compound layer grafted on the CIMS surface\nintroduces abundant amino groups, imparting positive charges to the membrane\nsurface. This positively charged layer strongly repels multivalent cations such as Ni²⁺,\nCo²⁺, and Mn²⁺, while monovalent cations (e.g., Na⁺, Li⁺) with lower charge density\nexperience much weaker repulsion, enabling preferential transport (Fig. S5a). In\naddition, the relatively dense structure of the CIMS membrane matrix provides a size\nsieving effect, where Na⁺ and Li⁺, with smaller hydrated radii, permeate more easily\nthan the larger hydrated Ni2+, Co2+, and Mn2+ (Fig. S5b). The synergy between\nelectrostatic repulsion and size sieving enhances monovalent selectivity, thereby\ncontributing to the optimized design and operational efficiency of the dual-stage\nselective electrodialysis system. The potential mechanism could be the chemical\naffinity, which relates to chemical interactions between molecules, ions, or compounds.\nIn membrane separation processes, the chemical affinity is the tendency of molecules,\nions, or compounds to chemically interact with active groups within the membrane or\non the membrane surface. The most important chemical interaction between these\nspecies is bonding and is described by bonding affinity. As shown in Fig. S5c, under\nthe influence of a driving force, molecules, ions, or compounds with higher\nchemical/binding affinity (f ) to the membrane structure will more readily interact with\nca\nthe membrane surface and transport across the membrane. In contrast, ions with a\nhigher mass transfer resistance will not be able to cross the membrane surface and\nare selectively retained.\nSupplementary Figure S5. The mechanisms of commercial monovalent ion-selective\nelectro-membranes. (a) Electrostatic repulsion; (b) Size exclusion; (c) Chemical affinity.\nMore detail information had been published in our previous work:\nZhao, Y. et al. Advanced ion transfer materials in electro-driven membrane processes\nfor sustainable ion-resource extraction and recovery. Prog. Mater. Sci. 128, 100958\n(2022).\n2.3. Parameter equations in the manuscript were duplicated in those in the\nsupplementary information, such as the product recovery rate, the total suspended\nsolids measurement, the specific energy consumption.\nResponse:\nWe appreciate the reviewer’s observation and have deleted the redundant equations\nfrom the Supplementary Information, including those for product recovery rate, total\nsuspended solids measurement, and specific energy consumption.\n2.4. The e-MCGR process configurations (Lines 418-503) overlap significantly with\nthe membrane stack design and operating/testing descriptions in the supplementary\ninformation (Lines 111-193). Redundancy should be minimized.\nResponse:\nWe appreciate the reviewer’s observation regarding the redundancy. Accordingly, we\nhave removed the membrane stack design and operating/testing descriptions from the\nSupplementary Information to avoid overlap with the e-MCGR process configurations\npresented in the main manuscript.\n3. As depicted in FIGURE 4b, the Li CO crystal formed in the base stream loop.\n2 3\nHowever, the method and mechanism to prevent the deposition of Li CO crystals\n2 3\nwithin the membrane stack remain unexplained. Besides, for the in-situ crystallization\nmodule, the control strategy of the CO inlet flow rate should be clarified. It is\n2\nrecommended to provide a complete device photo.\nResponse:\nThanks for your comments. (1) In our current laboratory-scale setup, we observed that\nLi CO crystals on the membrane surface can be easily cleaned with deionized water\n2 3\nafter the experiment, so no specific treatment was applied to the deposited Li CO\n2 3\nduring the process. For the future scale-up experiments, two measures can be\nadopted. The first one is to perform aeration treatment on the obtained base solution\nafter the experiment, which allows Li CO crystals to suspend in the base chamber,\n2 3\nthereby avoiding crystallization inside the membrane stack. The second method is to\nconnect the BMED system with the ultrafiltration membrane module. A module\ncombining an ultrafiltration membrane with BMED has been designed. In the base\nchamber, a precisely controlled CO aeration system induces Li CO crystallization.\n2 2 3\nThe BMED effluent, containing suspended Li CO crystals, flows to a cross-flow\n2 3\nultrafiltration module (UP150 membrane). The ultrafiltration permeate is then\nrecirculated to the BMED stack. The Li CO crystals can be recovered on the\n2 3\nultrafiltration membrane surface.\n(2) For the in-situ crystallization module, the CO inlet flow rate is monitored and\n2\ncontrolled using a flow meter. Fig. S15 presents a photo of our complete device, which\nincludes the in-situ crystallization module and related components. The CO was\n2\ncontinuously introduced into the base solution through a pipe at the bottom of the 2 L\nreservoir. A constant rate of 50 mL/min through a mass flow controller.\nSupplementary Figure S15. The complete device photo of the bipolar membrane in-\nsitu crystallization system.\n4. In the line 456 of the manuscript and the line 151 of the supplementary information,\nthe Ksp of Li CO should be about 0.0016 instead of 10^(-10.33).\n2 3\nResponse:\nWe thank the reviewer for pointing out the inconsistency regarding the solubility of\nLi₂CO₃. We have checked it again, and the Ksp of Li CO should be 8.15×10−4\n2 3\naccording to CRC Handbook of Chemistry and Physics, 84th edition (2004), which is\nequivalent to a molar solubility of 0.0597 mol/L or 0.44 g/100 mL. We have now\nupdated the manuscript and supplementary information and have used the 8.15×10−4\nwhere appropriate for discussing crystallization.\n5. FIGURE S9 shows no permselectivity between Li+ and Ni2+ could not be observed,\nwhile the monovalent cation exchange membrane could reject most Ni2+ in Fig 3 and\nFig S8, which is contradictory. The experimental results reflected in Figure S9 should\nbe checked carefully, and the corresponding textual descriptions should be revised to\naid readers in comprehending the ideas the authors intend to convey.\nResponse:\nWe are sorry for the confusion. We have removed Fig. S9. The Fig. S9 here shows\nthe Li+ and Ni2+ permeation in a commercial cation exchange membrane that has no\nmonovalent selectivity. Initially, we wanted to compare Fig. S9 of CEM with no\nselectivity to Fig. 3 of CIMS, to state the term of ion-transfer kinetics (e-MK). However,\nafter careful consideration, we decided to delete Fig. S9 as it is not necessary to prove\nno monovalent selectivity of CEM and monovalent selectivity of CIMS.\n6. In line 214-216, the phenomenon of mineral scaling on the membranes at high\ncurrent density of 5 mA/cm2 was attributed to the poor anti-scaling properties of\nexisting membranes. This explanation is unreasonable. As shown in FIGURE S12, the\nscaling phenomenon on the surface of the CIMS membrane was more severe than\nthat of the AMX membrane. This should be attributed to the occurrence of water\nsplitting on the surface of CIMS. Due to that most divalent ions are repelled by CIMS\nand are unable to serve as current carriers, water would be splitted as H+ and OH- to\nsupplement the lack of current carrier flux at the cross section of CIMS, which would\nbe more obvious at high current density.\nResponse:\nThanks for your critical observation. We agree that the previous attribution to the poor\nanti-scaling properties of the membranes was incomplete. Due to the polyamino\ncompound layer on the CIMS surface enhances electrostatic repulsion, an ion\nconcentration gradient forms in the boundary layer. Although the operating current\ndensity was below the measured limiting current density, the reduced ion\nconcentration in the boundary layer will elevate ion depletion, creating a localized\nmicro-environment where ion transport is imbalanced. Consequently, water molecules\nin the boundary layer undergo forced ionization to generate H⁺ and OH⁻ for charge\ncompensation. The OH- reacts with accumulated metal ions to form insoluble\nhydroxide precursors for membrane scaling formation (Fig. S13).\nWe revised the explanation of scaling in the manuscript and provided a schematic\nillustration of the scaling phenomenon in the supplementary materials:\n“While selective membrane dual-stage distillation technology achieves a high\nseparation efficiency, current commercial membranes are limited by their metal ion\nthroughput. In addition, they suffer from severe membrane scaling (Fig. S11 and Fig.\nS12). On the one hand, this scaling can be attributed to the inherent poor anti-scaling\nproperties of commercial electro-membrane. On the other hand, water splitting may\noccur on the CIMS surface as most divalent ions are electrostatically repelled and\ncannot serve as current carriers. The resulting shortage of charge carriers induces\nforced dissociation of water into H⁺ and OH⁻, with OH⁻ reacting with accumulated\nmetal ions to form insoluble hydroxide precursors that deposit on the membrane\nsurface (Fig. S13).”\nSupplementary Figure S11. The pictures of membranes before operations.\nSupplementary Figure S12. The pictures of membranes after operations (under the\ncurrent density of 5.0 mA cm−2) show mineral scaling on the membrane surface.\nSupplementary Figure S13. The concentration polarization-induced scaling on the\nCIMS membrane.\n7. The electro-membrane processes operated at constant current densities. However,\nneither the manuscript nor the supplementary information reflected any voltage curves\nof membrane stacks during the operation. These curves are crucial for assessing\nprocess stability and concentration polarization. Relevant data and discussion must\nbe added.\nResponse:\nThanks for your suggestions. Relevant data and discussion have been added in Fig.\nS9 in the Supplementary Information.\nFig. S9 illustrates the voltage drop changes for the multi-stage selective electrodialysis\noperations. The voltage drop across the selective electrodialysis stack changes\nsmoothly during the operation, which indicates the constant migration of ions through\nion exchange membranes at different stages. The voltage drop at the ends of both\nselective electrodialysis stacks increased with an increase in current density, followed\nby a decrease and stabilization (Figs. S9b-c). According to Ohm’s law, at a constant\ncurrent density, higher resistance leads to a higher voltage. At the beginning of the\nexperiment, ions entering the concentration chamber through the CIMS membrane\nand AMX membrane cause the resistance of the membrane stack to decrease.\nSubsequently, ions migrate steadily to adjacent chambers, keeping the resistance of\nthe membrane stack almost constant. The passage of Li⁺ and blocking of Ni²⁺, Co²⁺,\nand Mn²⁺ occur gradually under the action of the electric field, resulting in the voltage\ndrop first decreasing and then stabilizing.\nIn the bipolar membrane in-situ crystallization system, the voltage drop initially\nexhibited a higher value due to the elevated resistance caused by deionized water in\nthe acid/base chambers (Fig. S9d). Subsequently, the bipolar membrane generated\nH⁺/OH⁻, which combined with anions/cations from the feed chamber to form acid/base,\nleading to a stable voltage trend. In the later stage of the experiment, a large number\nof ions were transferred from the feed chamber, causing the resistance of the feed\nchamber to become a major contributor to the overall stack resistance; consequently,\nthe membrane stack resistance increased rapidly. As a result, the voltage of the\nmembrane stack rose sharply in the later stages of the experiment.\nIn complexation electrodialysis, the chelated species formed after chelation with\ndivalent transition metal ions in the leachate exist as anionic species ([EDTA-Co]²⁻\nand [EDTA-Ni]²⁻) (Fig. S9e). Similarly, the voltage drop initially decreases, then\nstabilizes, and finally increases. At the beginning of the experiment, [EDTA-Co]²⁻ and\n[EDTA-Ni]²⁻ ions migrate into the complexing chamber through the CMX membrane,\nwhile Mn²⁺ ions enter the recovery chamber through the AMX membrane; these\nprocesses collectively lead to a decrease in the membrane stack resistance.\nSubsequently, ions migrate steadily into adjacent chambers, maintaining the\nmembrane stack resistance at an almost constant level. In the later stage of the\nexperiment, the depletion of ions in the feed chamber causes its resistance to increase\nrapidly.\nIn the membrane-based metal-extracting temporal crystallization system, the voltage\ninitially exhibited a higher value due to the presence of deionized water in the Ni-\nrecovery chamber at the start of the experiment. Under the applied electric field, Ni²⁺\nand Co²⁺ ions then migrated into the extraction chamber through the CMX membrane,\nleading to a decrease in the membrane stack resistance (Fig. S9f). Within this process,\nCo²⁺ was extracted by Cyanex 272, while Ni²⁺ continued to migrate into the Ni-recovery\nchamber through the CMX membrane. Subsequently, ions transferred steadily to\nadjacent chambers, maintaining the membrane stack resistance at an almost constant\nlevel. Finally, as the ion concentration in the feed chamber became low, the voltage of\nthe system showed an increasing trend.\nSupplementary Figure S9. (a) Schematic of the power supply for voltage drop\nmeasurement. Voltage drop versus time curves of (b) first-stage selective\nelectrodialysis, (c) second-stage selective electrodialysis, (d) bipolar membrane in-situ\ncrystallization system, (e) complexation electrodialysis, and (f) membrane metal-\nextracting temporal crystallization system.\n8. The influence of the molar ratio of EDTA to Ni2+ and Co2+ was not discussed in the\nmanuscript. However, the molar ratio is a key parameter determining the speciation of\nNi2+, Co2+ and Mn2+ in the wastewater. A lower molar ratio is not conducive to the full\nconversion of Ni2+ and Co2+ into [EDTA-Ni]2- and [EDTA-Co]2- forms, which in turn is\nnot conducive to the recovery of Ni2+ and Co2+, and will result in the recovered Mn2+\nconcentrate containing Ni2+ and Co2+ impurities. An excessively high molar ratio will\nturn part of Mn2+ into [EDTA-Mn]2- form, thereby reducing the recovery rate of Mn2+\nand causing the recovered Ni2+ and Co2+ concentrates to contain Mn2+ impurities.\nTherefore, relevant results and discussion should be supplemented.\nResponse:\nThanks for your comments and suggestions. We have added details in the manuscript.\nThe influence of the EDTA: (Ni+Co) molar ratio, including 0.95, 1.0, 1.05, and 1.15, on\nmetal ion separation is explored (Fig. S26). The concentration of Ni²⁺, Co²⁺ and Mn²⁺\nin the feed chamber presented a decreasing trend (Figs. S26a-c). The Ni²⁺ and Co²⁺\nform Ni[EDTA]2+/Co[EDTA]2+ complex anions, thus they migrate towards the anode\ninto the chamber. On the contrary, the Mn2+ migrated towards the cathode into the\nrecovery chamber as the experiment progressed (Fig. S26f). A small amount of Ni2+ and\nCo2+ is also recovered along with Ni2+ when the molar ratio of EDTA to Ni+Co at 1,\ndue to the incomplete complexation (Figs. S26d-e). As the molar ratio of EDTA to\nNi+Co increases to 1.05, no extra Ni2+ and Co2+ ions are observed in the recovery\nchamber, and Mn2+ approaches the highest recovery. However, a further increase in\nthe molar ratio to 1.15 significantly decreased the recovery efficiency of Mn2+ due to\nthe formation of Mn[EDTA]2+. Considering the higher recovery of Mn2+, the molar ratio\nof EDTA to Ni+Co is optimized at 1.05. With the extra-added H+, which is supplied from\nthe bipolar membrane in-situ crystallization, [EDTA-Ni]2- and [EDTA-Co]2- de-coordinate with\nEDTA, and the subsequent separation is carried out in the membrane metal-extracting\ntemporal crystallization configuration.\nBesides, we have discussed more details in the Supplementary Information. In Fig.\nS26, different molar ratios of EDTA to Ni+Co, i.e., 0.95, 1.0, 1.05, 1.1, and 1.15, are\nstudied to increase the recovery efficiency of Mn2+ ions.\n1.2. Influence of the molar ratio of EDTA to Ni2+ and Co2+\nA range of molar ratios of EDTA-4Na to Ni2+/Co2+ (i.e., 0.95, 1.0, 1.05, 1.1, and 1.15)\nis tested to maximize Mn2+ recovery efficiency in the recovery chamber. The current\ndensity of the complexation electrodialysis stack is 5 mA cm-2.\nSupplementary Figure S26. Effect of the molar ratio of EDTA to Ni+Co on the ion\nconcentration in (a, b, c) feed chamber, (d, e, f) recovery chamber, and (g, h, i)\ncomplexation chamber.\n9. In FIGURE 5h, the horizontal axes are labeled as \"Feed chamber\", \"Feed chamber\",\nand \"Recovery chamber\", respectively. However, according to FIGURE S25, they\nshould be \"Feed chamber\", \"Co-extraction chamber\", and \"Recovery chamber\",\nrespectively. Meanwhile, the concentration evolution of Ni2+ and Co2+ in Co-extraction\nchamber and Recovery chamber was confusing. Within 0 to 10 hours, the\nconcentration of Co2+ in the Co-extraction chamber continued to rise, while the\nconcentration of Co2+ in the Recovery chamber remained at a low level. This was\nmainly due to the complexion of Co2+ to Cyanex272 formed the non charged\nCo(Cyanex272) . After 10 h, the Co2+ concentration in the Co-extraction chamber\n2\ndeclined while that in recovery chamber increased with time. It seemed that H SO\n2 4\nwas added into the feed solution to substitute Co2+ from Co(Cyanex272) (lines 184-\n2\n187 in supplementary information) at the operation time of 10h, but the manuscript did\nnot explicitly clarify this point.\nResponse:\nWe thank the reviewer for pointing out this. In response, (1) we have revised the label\nof the horizontal axes in Fig. 5h to feed chamber, co-extraction chamber and recovery\nchamber.\nFig. 5. Recycling of Mn2+, Ni2+, and Co2+ from the divalent cation solutions. a, Schematic\nof the subsequent membrane metal-complexing ex-situ crystallization as well as membrane\nmetal-extracting temporal crystallization configurations for sequential separation of Mn2+, Ni2+,\nand Co2+; b, Mechanisms of the EDTA complexion (selective coordination with Ni2+ and Co2+)\nand the Cyanex 272 extraction (preferential binding to Co2+), enabling targeted metal recovery;\nc, Photographs of recovered products, (i) black Mn O , (ii) green Ni(OH) , and (iii) pink\n3 4 2\nCo(OH) powders; d, SEM images and elemental mapping of Mn O , Ni(OH) and Co(OH)\n2 3 4 2 2\ncrystals, respectively, confirming uniform morphology and homogeneous distribution of Mn,\nNi, and Co; e, XPS spectra showing atomic compositions of (i) Mn O (Mn: 26.9%, O: 37.6%),\n3 4\n(ii) Ni(OH) (Ni: 29.1%, O: 56.3%), and (iii) Co(OH) (Co: 23.7%, O: 54.9%), with <0.1%\n2 2\nimpurities; f, XRD spectra confirming phase-pure Mn O (PDF#24-0734), Ni(OH) (PDF#14-\n3 4 2\n0117), and Co(OH) (PDF#30-0443) powders, with no impurity peaks; g, Temporal evolution\n2\nof Mn2+, Ni2+, and Co2+ concentrations in membrane metal-complexing ex-situ crystallization\nsystem, achieving > 99.5% Mn2+ at 5 mA cm-2 and at EDTA : (Ni + Co) of 1.05; h, Temporal\nevolution of Ni2+ and Co2+ concentrations in membrane metal-extracting temporal\ncrystallization system, recovering > 83.1% Ni2+ and > 87.3% Co2+; i, Recovery and product\npurity of Mn O (>99.5% recovery, >99.9% purity), Ni(OH) (>83.1% recovery, >99.5% purity),\n3 4 2\nand Co(OH) (>87.3% recovery, >92.5% purity). j, The energy, economic, and environmental\n2\nanalysis: energy consumption (Mn O : 1.30 $ kg-1 Mn O ; Ni(OH) and Co(OH) : 1.89 $ kg-1\n3 4 3 4 2 2\n[Ni(OH) + Co(OH) ]), investment costs (Mn O : 1.02 $ kg-1 Mn O ; Ni(OH) and Co(OH) : 1.86\n2 2 3 4 3 4 2 2\n$ kg-1 [Ni(OH) + Co(OH) ]), and net profits (Mn O : 6.68 $ kg-1 Mn O ; Ni(OH) and Co(OH) :\n2 2 3 4 3 4 2 2\n12.37 $ kg-1 [Ni(OH) + Co(OH) ]).\n2 2\n(2) We have discussed more details in the manuscript about the Co2+ changes: \"The\ntemporal evolution of Ni2+, Co2+, and Mn2+ concentrations in each chamber is\nmeasured using inductively coupled plasma optical emission spectrometry (ICP-OES).\nIn the membrane metal-complexing ex-situ crystallization configuration (Fig. 5g), the\nNi2+ and Co2+ form [EDTA-Ni]2- and [EDTA-Co]2- and migrate to the Ni & Co collect\nchamber in the anode side, while Mn2+, which remains uncomplexed, transports\ntowards the cathode into the Mn recovery chamber. Then [EDTA-Ni]2- and [EDTA-\nCo]2- are decomplexed using 2.0 M H SO , the collected Ni and Co solution is used in\n2 4\nthe next membrane metal-extracting temporal crystallization configuration (Fig. 5h). In\nthe first 10 h, the Co2+ accumulates in the co-extraction chamber as the Co2+ \ncomplexes with Cyanex272, forming non-charged Co(Cyanex272) , while positively\n2\ncharged Ni2+ migrates to the recovery chamber. After 10 h, Co(Cyanex272) is\n2\ndissociated into Co2+, which transports to the recovery chamber later.”\n10. FIGURE 5i indicates that the Ni2+ recovery rate at 10 h was lower than 80%, but\nauthors claimed that the recovery rate could reach 92.5%. Please provide a\nreasonable explanation.\nResponse:\nWe apologize for the oversight. The Ni²⁺ recovery rate should be corrected to 83.1%\n(at 12 h), rather than as previously stated. We have revised the manuscript accordingly.\nFig. 5i, Recovery and product purity of Mn O (> 99.5% recovery, > 99.9% purity), Ni(OH) (>\n3 4 2\n83.1% recovery, > 99.5% purity), and Co(OH) (> 87.3% recovery, > 92.5% purity).\n2\n11. No feasible method for recycling EDTA and Cyanex 272 was presented or\nvalidated. Given their cost and environmental impact, reagent recycling is crucial for\nthe feasibility of divalent ion separation methods.\nResponse:\nThanks for your comments. In our study, we developed a regeneration approach for\nboth EDTA and Cyanex 272 using the acid/base byproducts generated from the\nbipolar membrane in-situ crystallization system. In the proposed system, the HCl and\nNaOH can be recovered in the acid chamber and base chamber, respectively. Thus,\nwe have added related information in the Supplementary Information.\nMethod for recycling EDTA and Cyanex 272. In the e-MCGR system, HCl and\nNaOH can be recovered in the acid chamber and base chamber, respectively. The\nHCl concentration and NaOH concentration approached 1.03 mol/L and 1.1 mol/L,\nrespectively. After the membrane metal-complexing ex-situ crystallization, Ni and Co\nwere decomplexed from EDTA at pH below 1.5 by adding HCl. The EDTA solid was\nprecipitated out and recovered according to the following reaction:\n[𝐸𝐷𝑇𝐴−𝑁𝑖]2−+2𝐻+→ 𝑁𝑖2++2𝐸𝐷𝑇𝐴 ↓\n[𝐸𝐷𝑇𝐴−𝐶𝑜]2−+2𝐻+→ 𝐶𝑜2++2𝐸𝐷𝑇𝐴 ↓\nAfter the membrane metal-extracting temporal crystallization, the Cyanex 272 was\nregenerated by adding HCl at pH below 1 according to the following reaction:\n2𝐻++𝐶𝑜(𝐶𝑦𝑎𝑛𝑒𝑥272) →2𝐶𝑦𝑎𝑛𝑒𝑥272+𝐶𝑜2+\n2\n2𝐻++𝑁𝑖(𝐶𝑦𝑎𝑛𝑒𝑥272) →2𝐶𝑦𝑎𝑛𝑒𝑥272+𝑁𝑖2+\nrepeatedly recycled EDTA and Cyanex 272.\n12. A puzzling phenomenon in economic analysis part is that, as a high-value\nchemical, Li CO was less profitable than nickel and cobalt compounds in Table S7\n2 3\nand Table S9. Also, the cost of EDTA and Cyanex 272 was ignored in the economic\nanalysis. Authors should clarify these issues and supplement the comparison with\nmainstream technologies for lithium recovery.\nResponse:\nThanks for your comments. Upon re-evaluating the market data, we identified that our\ninitial assessment inadvertently used the price for analytical-grade Li CO rather than\n2 3\nthe higher-value battery-grade product that matches our specifications. In the revised\nmanuscript, we added a new market analysis of the battery-grade Li CO product (24\n2 3\n$ kg-1). The revision better reflects the commercial value of our high-purity output while\nmaintaining conservative estimates for nickel/cobalt byproducts.\nIn the revised manuscript, we have now fully incorporated the costs of EDTA and\nCyanex 272 into our economic analysis (Membrane and chemicals cost section). The\nmodified economic analysis is shown in the following Tables.\nWe would like to clarify that while mainstream technologies (e.g., chemical\nprecipitation, solvent extraction, and adsorption) can effectively remove metal ions\nfrom the spent LIBs stream. However, these technologies typically focus solely on ion\nremoval without enabling the recovery of metal ions as valuable products. In contrast,\nour technology achieves efficient separation of metal ions as high-purity products (e.g.,\nLi⁺, Ni²⁺, Co²⁺, Mn²⁺) with commercial value. Compared to the mainstream\ntechnologies for lithium recovery, the e-MCGR technology yields a profit of around\n20.79 $ kg-1 Li CO , 6.63 $ kg-1 Mn O , and 12.37 $ kg-1 for Ni(OH) + Co(OH) .\n2 3 3 4 2 2\nThis indicates that e-MCGR technology not only achieves a lower overall cost but also\noffers distinct profit advantages.\nSupplementary Table S11. The economic analysis of Li CO recovery.\n2 3\nOperating conditions Li CO\n2 3\nCurrent density (mA cm-2) 5\nEach ion-exchange membrane area (cm2) 297\nEnergy cost\nEnergy consumption of Li CO recovery (kg kWh-1) 0.095\n2 3\nProcess capacity per year (kg Y-1) 41.34\nElectricity price ($ kWh-1) 0.10\nEnergy cost for resource recovery ($ kg-1) 1.05\nEnergy cost for peripheral equipment ($ kg-1) 0.052\nTotal energy cost ($ kg-1) 1.10\nInvestment cost\nEntire device lifetime (Y)a 3\nMembrane cost ($) 70.06\nMembrane equipment cost ($)b 105.09\nAuxiliary equipment cost ($)c 73.57\nTotal investment cost ($) 178.66\nAmortization per year ($ Y-1) 69.33\nMaintenance cost per year ($ Y-1) d 17.87\nFixed cost per year ($ Y-1) 87.19\nFixed cost for Li CO recovery ($ kg-1) 2.11\n2 3\nOverall Li CO cost ($ kg-1) e 3.21\n2 3\nLi CO market prize ($ kg-1) f 24.00\n2 3\nTota Li CO profit ($ kg-1) 20.79\n2 3\na The lifetime of the experimental device was set as three years.\nb Membrane equipment cost equals 1.5 times the membrane cost.\nc Auxiliary equipment equals 1.5 times the stack cost.\nd Maintenance equals 10% of total investment cost.\ne The overall recovery cost included both energy cost and investment cost.\nf Data collected from Alibaba.com.\nSupplementary Table S12. The economic analysis of Mn O recovery.\n3 4\nOperating conditions\nEDTA/Ni+Co molar ratio 0.95 1 1.05 1.10 1.15\nEach ion-exchange membrane\n297 297 297 297 297\narea (cm2)\nEnergy cost\nEnergy consumption of Mn O 0.096 0.083 0.081 0.074 0.068\n3 4\nrecovery (kg kWh-1)\nProcess capacity per year (kg Y-1) 52.59 52.54 51.24 46.86 43.05\nElectricity price ($ kWh-1) 0.10 0.10 0.10 0.10 0.10\nEnergy cost for resource recovery 1.04 1.21 1.24 1.35 1.47\n($ kg-1)\nEnergy cost for peripheral 0.052 0.060 0.062 0.068 0.074\nequipment ($ kg-1)\nTotal energy cost ($ kg-1) 1.091 1.27 1.30 1.42 1.55\nInvestment cost\nEntire device lifetime (Y) a 3 3 3 3 3\nMembrane and chemicals cost ($) 42.03 42.05 42.07 42.10 42.12\nMembrane equipment cost ($)b 62.04 62.07 63.11 63.14 63.18\nAuxiliary equipment cost ($)c 44.13 44.15 44.18 44.20 44.23\nTotal investment cost ($) 107.17 107.23 107.29 107.35 107.41\nAmortization per year ($ Y-1) 41.58 41.61 41.63 41.65 41.68\nMaintenance cost per year ($ Y-1) d 10.72 10.72 10.73 10.73 10.74\nFixed cost per year ($ Y-1) 52.30 52.33 52.36 51.39 52.42\nFixed cost for Mn O recovery 0.99 1.00 1.02 1.12 1.22\n3 4\n($ kg-1)\nOverall Mn O cost ($ kg-1) e 2.09 2.26 2.32 2.54 2.76\n3 4\nMn O market prize ($ kg-1) f 9.00 9.00 9.00 9.00 9.00\n3 4\nTota Mn O profit ($ kg-1) 6.91 6.74 6.68 6.46 6.24\n3 4\na The lifetime of the experimental device was set as three years.\nb Membrane equipment cost equals 1.5 times the membrane cost.\nc Auxiliary equipment equals 1.5 times the stack cost.\nd Maintenance equals 10% of the total investment cost.\ne The overall recovery cost included both energy cost and investment cost.\nf Data collected from Alibaba.com.\nSupplementary Table S13. The economic analysis of Ni(OH) /Co(OH) recovery.\n2 2\nOperating conditions\nCurrent density (mA cm-2) 3 5 7\nEach ion-exchange membrane area (cm2) 297 297 297\nProportion of Ni2+ (%) 31.14 46.85 33.53\nProportion of Co2+ (%) 68.86 53.15 66.47\nEnergy cost\nEnergy consumption (kg kWh-1) 0.075 0.056 0.057\nProcess capacity per year (kg Y-1) 26.57 33.43 27.52\nElectricity price ($ kWh-1) 0.10 0.1 0.1\nEnergy cost for ($ kg-1) 1.34 1.80 1.76\nEnergy cost for peripheral equipment ($ kg-1) 0.067 0.090 0.088\nTotal energy cost ($ kg-1) 1.41 1.89 1.85\nInvestment cost\nEntire device lifetime (Y) a 3 3 3\nMembrane and chemicals cost ($) 50.00 50.00 50.00\nMembrane equipment cost ($)b 75.00 75.00 75.00\nAuxiliary equipment cost ($)c 52.50 52.50 52.50\nTotal investment cost ($) 127.49 127.49 127.49\nAmortization per year ($ Y-1) 49.47 49.47 49.47\nMaintenance cost per year ($ Y-1) d 12.75 12.75 12.75\nFixed cost per year ($ Y-1) 62.22 62.22 62.22\nFixed cost for Ni(OH) /Co(OH) recovery 2.34 1.86 2.26\n2 2\n($ kg-1)\nOverall Ni(OH) /Co(OH) cost ($ kg-1) e 3.75 3.75 4.11\n2 2\nNi(OH) market prize ($ kg-1) f 14.00 14.00 14.00\n2 2\na The lifetime of the experimental device was set as three years.\nb Membrane equipment cost equals 1.5 times the membrane cost.\nc Auxiliary equipment equals 1.5 times the stack cost.\nd Maintenance equals 10% of total investment cost.\ne The overall recovery cost included both energy cost and investment cost.\nf Data collected from Alibaba.com.\nSupplementary Table S14. Comparison of e-MCGR with mainstream technologies.\nMethods Cost ($ kg-1) References\nHydrometallurgical 27.7 18\nElectrodialysis 6.34 19\nChemical extraction 127.6 20\nAdsorption 13.05 21\nExtraction and precipitation 52.79 22\nSelective leaching 6.19 23\n18. Cao, Y. et al. Co-products recovery does not necessarily mitigate environmental\nand economic tradeoffs in lithium-ion battery recycling. Resour., Conserv. Recycl. 188,\n106689 (2023).\n19. Shan, W. et al. Producing battery grade lithium carbonate from salt‐lake brine via\nbipolar membrane carbon dioxide mineralization. AIChE J. 71, e18675 (2025).\n20. Ali-Zada, Z., Şengil, İ.A. & Alp, A. Environmentally friendly and economic method\nfor recovery of metals from spent lithium-ion batteries using\nethylenediaminetetraacetic acid disodium salt and citric acid. Waste Manage. 204,\n114896 (2025).\n21. Lin, X. et al. Selective recovery of Ni, Co, and Li from spent NCA lithium-ion\nbatteries: 3D macroporous bio-based adsorbent and innovative leaching–adsorption\nstrategy. J. Environ. Chem. Eng., 118063 (2025).\n22. Li, C. et al. Separation and recovery of nickel cobalt manganese lithium from waste\nternary lithium-ion batteries. Sep. Purif. Technol. 306, 122559 (2023).\n23. Kang, X. et al. Selective recovery of lithium from used lithium-ion batteries spent\nvia grain boundary reconstruction reaction. Sep. Purif. Technol. 356, 129726 (2025).\nWe also have recalculated the energy, economic and environmental analysis of the\nsystems of bipolar membrane in-situ crystallization, membrane metal-complexing ex-\nsitu crystallization and membrane metal-extracting temporal crystallization. We have\nrevised them in Fig. 4h and Fig. 5j.\nFig. 4h, The energy, economic and environmental analysis of the bipolar membrane\nin-situ crystallization system: energy cost (1.10 $ kg-1 Li CO ), investment (2.11 $ kg-\n2 3\n2 3 2 3 2 3\n$ kg-1).\nFig. 5j, The energy, economic, and environmental analysis: energy consumption\n(Mn O : 1.30 $ kg-1 Mn O ; Ni(OH) and Co(OH) : 1.89 $ kg-1 [Ni(OH) + Co(OH) ]), \n3 4 3 4 2 2 2 2\ninvestment costs (Mn O : 1.02 $ kg-1 Mn O ; Ni(OH) and Co(OH) : 1.86 $ kg-1\n3 4 3 4 2 2\n[Ni(OH) + Co(OH) ]), and net profits (Mn O : 6.68 $ kg-1 Mn O ; Ni(OH) and Co(OH) :\n2 2 3 4 3 4 2 2\n12.37 $ kg-1 [Ni(OH) + Co(OH) ]).\n2 2\n13. In line 226-268 of the supplementary information, the methods of environmental\nanalysis, economic analysis and energy analysis were described vaguely, without\nnecessary data and evaluation methods to support these analysis.\nResponse: Thanks for your comment. To address this, we provide the detailed\nmethodologies for environmental analysis, economic analysis, and energy analysis in\nthe supplementary materials. Moreover, to make this analysis more robust and clear\nin the revised manuscript, we will supplement specific data to support these claims.\n(1) Economic analysis: An economic analysis is important for assessing the potential of the\nintegrated system for achieving near zero discharge from spent LIBs wastewater. The overall\nrecovery cost for spent LIBs wastewater contained energy cost and investment cost. The\nlifetime of the integrated system device was set as three years and relied on lab-scale data.\nIn pilot plant tests and industrial-scale operations, fixed costs were amortized as production\nvolume increased, reducing unit costs and enhancing economic attractiveness. The optimized\ne-MCGR process achieves a recovery of 95.5% for Li+, 99.5% for Mn2+, 83.1% for Ni2+, and\n87.3% for Co2+, yielding high product purities of 99.9% for Li CO , 99.9% for Mn O , 99.5%\n2 3 3 4\nfor Ni(OH) and 92.5% for Co(OH) . Thus, the system yields a profit of around 20.79 $ kg-1\n2 2\nLi CO , 6.68 $ kg-1 Mn O , and 12.37 $ kg-1 for Ni(OH) + Co(OH) The increased operation\n2 3 3 4 2 2. \nscale also enhanced operational efficiency and by-product utilization, such as promoting\nLi CO , Ni(OH) , Co(OH) , and Mn O recovery and reducing energy consumption, which\n2 3 2 2 3 4\npositively impacted the economic viability of the project. The economic analysis is calculated\naccording to the method in Table S10.\n(2) Energy analysis: From an energy consumption perspective, although the\nelectrodialysis process requires electrical energy, the overall energy balance is\nfavorable when considering the value of the recovered metals. When evaluating the\noverall energy balance, it is essential to consider the energy savings achieved by\navoiding primary metal mining. The extraction of lithium, nickel, cobalt, and\nmanganese from the natural ores is an extremely energy-intensive process. It involves\nactivities such as ore mining, crushing, grinding, and complex chemical separation\nprocesses. By recycling these metals from spent LIBs wastewater through integrated\nelectrodialysis technology, a significant amount of energy consumed in primary mining\ncan be saved. In addition, the energy efficiency of the electrodialysis integrated system\ncan be further improved through process optimization. The adjustment of operating\nparameters, such as current density, to achieve a higher crystallization rate with the\nlowest energy input. The optimized e-MCGR process achieves an energy\nconsumption of 11.83 kWh kg-1 for the first-stage selective electrodialysis, an energy\nconsumption of 15.97 kWh kg-1 for the second-stage selective electrodialysis, an\nenergy consumption of 10.45 kWh kg-1 for the bipolar membrane in-situ crystallization\nsystem, and an energy consumption of 17.96 kWh kg-1 for the membrane metal-\ncomplexing ex-situ crystallization and membrane metal-extracting temporal\ncrystallization. Moreover, the possible integration of renewable energy sources will be\nuseful. For example, solar or wind power can be used to supply part of the electrical\nenergy required for the integrated electrodialysis process, which not only reduces the\nreliance on traditional fossil fuel-based energy but also further improves the overall\nenergy sustainability of the metal recycling process. Thus, our method shows great\npotential in terms of energy efficiency and sustainability.\n(3) Environmental analysis: The environmental analysis in this study is conducted\nthrough a comparison between traditional processes (e.g., pyrometallurgy and\nhydrometallurgy) and the e-MCGR technology. This focuses on highlighting the\nenvironmental benefits of the e-MCGR technology of spent LIBs wastewater recovery.\nThe details are as follows: The implementation of electro-driven membrane technology\nsignificantly reduces the environmental burden associated with the disposal of spent\nLIBs wastewater. Traditional disposal methods (e.g., pyrometallurgy and\nhydrometallurgy) often lead to the release of heavy metals (such as nickel, cobalt, and\nmanganese) into the environment, which will cause soil and water pollution. However,\nextra reagent consumption, significant waste emission, secondary pollutant\ngeneration and complicated separation routes are limited. By recovering valuable\nelements (i.e., Ni2+, Co2+ and Mn2+) through electro-driven membrane technology, the\npotential for environmental contamination is greatly mitigated. This not only aligns with\nthe principles of the circular economy but also reduces the volume of waste that needs\nto be disposed of in landfills. The sustainable membrane-based technologies promote\nthe development of environmentally friendly strategies in the battery industry.\nSupplementary Table S10. The economic analysis method.\nCost types\nEnergy cost\nEnergy consumption of product recovery Collected from the energy consumption of the\n(kg kWh-1) membrane stack\nEffective running time of membrane module\nper year (365 days × 24 h/day × 0.8) ÷ time\nProcess capacity per year (kg Y-1) consumed by membrane module operation\n(h) × resource recovery by membrane stack\noperation (kg)\nElectricity price ($ kWh-1) 0.10\nEnergy cost for resource recovery ($ kg- Electricity price ($ kWh-1) ÷ Energy\n1) consumption of product recovery (kg kWh-1)\nEnergy cost for peripheral equipment Energy cost for resource recovery ($ kg-1) ÷\n($ kg-1) 20\nEnergy cost for resource recovery ($ kg-1) +\nTotal energy cost ($ kg-1)\nEnergy cost for peripheral equipment ($ kg-1)\nInvestment cost\nEntire device lifetime (Y)a 3\nArea of ion exchange membrane (m2) ×price\nMembrane cost ($)\nof ion exchange membrane ($ m-2).\nMembrane equipment cost ($) Membrane cost ($) × 1.5\nAuxiliary equipment cost ($) Membrane equipment cost ($) × 0.7\nMembrane cost ($) + Membrane equipment\nTotal investment cost ($)\ncost ($)\nTotal investment cost × 1.083-Amortization\nAmortization per year ($ Y-1) per year ($ Y-1) × 1.08 - Amortization per year\n($ Y-1) ×1.082\nMaintenance cost per year ($ Y-1) Total investment cost ($) ÷ 10\nAmortization per year ($ Y-1) + Maintenance\nFixed cost per year ($ Y-1)\ncost per year ($ Y-1)\nFixed cost per year ($ Y-1) ÷ Process capacity\nFixed cost for product recovery ($ kg-1)\nper year (kg Y-1)\nFixed cost for product recovery ($ kg-1) + Total\nOverall product cost ($ kg-1)\nenergy cost ($ kg-1)\nProduct market price ($ kg-1) Collected from Alibaba.com\nPrice of product ($ kg-1) - Overall product cost\nTotal product profit ($ kg-1)\n($ kg-1)\nIn this e-MCGR process, the specific energy consumption (E, kWh kg-1) is calculated\nby,\nt UId\nt t\nE=∫\nCVM\n0 t t\nwhere U (V) is the electro-membrane stack voltage, I (A) is the current, V (L) is the\nt t\nvolume of solution, C (mol L-1) is the concentration, and M (g mol-1) is the molecular\nt\nmass. More data on energy, economics and the environment on this e-MCGR\ntechnology are analyzed and summarized in the Supplementary Information.\n14. The FIGURE 1 in manuscript was not a core content and should not be included\nin the main text. Also, the cited literature was less relevant to the research content.\nResponse:\nWe appreciate the reviewer’s comment regarding Figure 1 and the cited literature. (1)\nWhile we acknowledge that Figure 1 is not part of the core experimental results, we\nrespectfully believe it serves an important purpose in establishing the broader context\nand motivation for this study. Specifically. It highlights the rapid growth in global LIBs\nproduction and consequent waste accumulation. By presenting these up-to-date data,\nthe figure underscores the pressing need for sustainable and efficient strategies to\nmanage end-of-life LIBs. We believe this contextual information helps readers better\nunderstand the real-world relevance and urgency of the recycling approach developed\nin our work.\n(2) The cited literatures are selected to highlight the growing environmental and\nresource challenges posed by the increasing volume of spent LIBs and the importance\nof developing efficient recycling technologies. This will ensure the citations more\neffectively bridge the significance of the spent LIBs recovery by using e-MCGR\ntechnology.\n15. In line 360-361, authors claimed that “the extra added NaOH and HCl are supplied\nand reused from the bipolar membrane in-situ crystallization system”. It lacked\nnecessary data support from a continuous integrated e-MCGR process.\nResponse: Thanks for your comments. In the revised manuscript, we provide the\nspecific data from the integrated e-MCGR process to validate the reuse of NaOH and\nHCl. In the bipolar membrane in-situ crystallization system, the HCl concentration and\nNaOH concentration approached 1.03 mol/L and 1.10 mol/L, respectively. After the\nmembrane metal-complexing ex-situ crystallization, 58 mL NaOH with a concentration\nof 1.10 mol/L was added to the solution for the precipitation of Mn O . After the\n3 4\nmembrane metal-extracting temporal crystallization, 55 mL NaOH with a concentration\nof 1.1 mol/L was added to the solution for the precipitation of Ni(OH) Then, 51 mL\n2.\nNaOH with a concentration of 1.10 mol/L was added to the solution for the precipitation\nof Co(OH) . We have mentioned it in Methods in manuscript:\nAll experimental configurations described herein were conducted using a single,\nconsistently prepared complex feed batch (the composition is shown in Supplementary\nTable S7) to ensure comparability and integration across the e-MCGR process. This\nbatch simulated real LIB leachate and was used sequentially across all experimental\nstages without intermediate reconstitution (Table S19). The integrated design of the\ne-MCGR process comprises:\ni. Selective membrane dual-stage distillation configuration: This configuration employs\ntwo sequential selective electro-membrane stacks to separate monovalent metal ions\n(Li+ and Na+) from the divalent metal ions (Mn2+, Ni2+ and Co2+). In the first-stage\nselective electro-membrane stack (Fig. S6), the membranes are arranged as: anode /\nbipolar membrane (BM) / AMX (commercial anion exchange membrane) × 3 / CIMS\n(commercially monovalent cation exchange membrane) × 3 / bipolar membrane (BM)\n/ cathode (Table S20). The effective area per membrane is 0.0189 m2. Four chambers\n(feed, transitional-1, transitional-2, recovery) are connected to containers holding 2 L\nsimulated LIB leaching solution in the feed chamber and 0.5 L deionized water in the\nother chambers. The operational parameters of the current densities are tested at 1\nmA cm-2, 2.5 mA cm-2, and 5 mA cm-2. The electric current of the electro-membrane\nstack is supplied by a stabilized power supply. In order to further distill Li+ and Na+\nfrom transitional chambers (transitional chamber-1 and transitional chamber-2) of the\nfirst stage, the second-stage selective electro-membrane stack is developed (Fig. S6),\nand the electro-membrane stack arrangement is anode / BM / AMX × 2 / CIMS × 2 /\nBM / cathode. Three chambers in the electro-membrane stack consist of the feed\nchamber, the transitional chamber, and the recovery chamber. For those chambers, 1\nL of residual solution from the first-stage transitional chambers is introduced into the\nfeed chamber of the second-stage stacks, and the transitional and recovery chambers\nare initially fed with 0.5 L of deionized water. The current density of the electro-\nmembrane stack is 0.5 mA cm-2. The refined Li+ and Na+ from the recovery chamber\nare then directly used as feed for the subsequent bipolar membrane in-situ\ncrystallization (ii) to separate Li+ from Na+.\nii. Bipolar membrane in-situ crystallization configuration: In this configuration, the\nelectro-membranes are stacked between anode and cathode in a sequence of the\nanode / BM / AMX / commercial cation exchange membrane (CMX) / BM / cathode\n(Fig. S13), forming three chambers, including the feed chamber, the acid chamber,\nand the base chamber. Specifically, 1 L monovalent metal ions solution (Li+ and Na+)\nfrom the recovery chambers of configuration (i) is used as the initial feed, and 0.5 L of\ndeionized water is fed to the acid and base chambers. The electric current of this\nelectro-membrane stack is 5 mA cm-2, which is supplied by a direct current-stabilized\npower supply. To avoid Na+ impurities in Li CO (caused by Na CO co-precipitation)\n2 3 2 3\nand capture the CO from industry, CO gas is introduced into the base chamber as a\n2 2\ncarbonate source, at a constant flow rate of 50 mL min-1 (Fig. S14). During the reaction,\nCO₂ dissolves into the solution and subsequently forms carbonate species. The Li CO\n2 3\ncrystals will be crystallized in the base solution, followed by Li CO recovery with the\n2 3\nUP150 membrane.\nCO + 2OH- → CO 2- + H O\n2 3 2\nThen, Li CO crystallization occurs as\n2 3\nCO 2- + 2Li+ → Li CO (K = 8.15×10-4)\n3 2 3 sp\nOverall, the reaction is\n2OH- + CO + 2Li+ →Li CO ↓+ H O\n2 2 3 2\niii. Membrane metal-complexing ex-situ crystallization configuration: Here, the\nmembrane sequence is: anode / CMX / AMX / CMX / AMX / cathode (Fig. S19), with\nthree chambers: the feed, complexing, and the recovery chamber. The residual\ndivalent metal-rich solution remaining after configuration (i) serves as the feed for this\nconfiguration. Herein, 3.0 L mixtures of Mn2+, Ni2+, and Co2+ are introduced into the\nfeed chamber, and 0.5 L of deionized water is fed into the complexing and recovery\nchambers. The feed solution is further treated with the EDTA-4Na salt in the feed\ncontainer. When the [EDTA]4− forms complex anions with Ni2+ and Co2+, with the\nEDTA’s differential coordination capacities (K > K >> K ), the Mn2+ is separated\nNi Co Mn\nfrom Ni2+ and Co2+ using this membrane metal-complexing ex-situ crystallization\nsystem,\n𝑁𝑖2++𝐸𝐷𝑇𝐴4− → [𝐸𝐷𝑇𝐴−𝑁𝑖]2− 𝐾 = 2.51×1018\n𝑁𝑖\n𝐶𝑜2++𝐸𝐷𝑇𝐴4− → [𝐸𝐷𝑇𝐴−𝐶𝑜]2− 𝐾 = 2.82×1016\n𝐶𝑜\n𝑀𝑛2++𝐸𝐷𝑇𝐴4− → [𝐸𝐷𝑇𝐴−𝑀𝑛]2− 𝐾 = 7.76×1013\n𝑀𝑛\nFor the decomplexation, EDTA complexes are dissociated by adding 2.0 M H SO\n2 4\nuntil pH < 0.5. To recover metal ions, Mn2+ in the recovery chamber is precipitated as\nMn(OH) by using 1.0 M NaOH. Both the acid and base used here are provided by\n2\nconfiguration (ii).\nMn2+ + 2OH- → Mn(OH) ↓\n2 3 4\n6Mn(OH) + O → 2Mn O ↓+ 6H O\n2 2 3 4 2\nMembrane metal-extracting temporal crystallization configuration: Connecting with the\nmembrane metal-complexing ex-situ crystallization configuration, this stack uses:\nanode / AMX / CMX / CMX / AMX / cathode (Fig. S19). In this system, the membrane\nstack contains three containers, which connect with the feed chamber, the Co-\nextraction chamber, and the Ni-recovery chamber. The containers are initially fed with\n0.5 L mixed Ni2+ and Co2+ solution from the membrane metal-complexing ex-situ\ncrystallization system, 0.5 L extraction solution (50% saponified Cyanex 272 in\nkerosene), and 0.5 L deionized water, respectively. The electric current of the\ncomplexation electrodialysis stack is tested at 5 mA cm-2. Due to the Cyanex 272\npreferentially binds Co2+, it enables efficient Ni2+ and Co2+ separation. In this process,\nthe Co2+ is extracted by Cyanex 272, which forms a stable complex with Cyanex 272\nin the extraction chamber.\nCo2++2Cyanex 272→Co(Cyanex 272) +2H+\nfield. The Ni2+ is precipitated as Ni(OH) with the usage of 1.0 M NaOH (supplied from\n2\nthe bipolar membrane in-situ crystallization system).\nNi2++ 2OH- → Ni(OH) ↓\n2 4\nrelease CO2+ from Cyanex 272,\nCo(Cyanex 272) +2H+→Co2++2Cyanex 272\nprecipitate as Co(OH) with the usage of 1.0 M NaOH.\nTo validate the practical feasibility of the e-MCGR process, real wastewater from spent\nternary LIBs is used as feed. The spent ternary LIBs, sourced from a technology\ncompany's power bank, are used as the primary raw materials. First, 0.4 g of waste\nNMC111 cathode material is leached in 4.0 M HCl at 80 ºC for 6 h. The resulting\nleachate is subjected to the same pretreatment sequence as described in Section\n“Pretreatment,” namely ultrafiltration (UP150 membrane, 0.01 MPa) to remove organic\nimpurities, followed by RO (BW30 membrane, 1.5 MPa) to concentrate the solution.\nThe concentrated leachate, now enriched in Li⁺, Na⁺, Mn²⁺, Ni²⁺, and Co²⁺, is directly\nused as the feed for the e-MCGR process, without reconstitution or modification, to\nensure a one-to-one comparison with the results obtained from the simulated LIB\nleachate."}, {"round": 2, "reviewer_comments": {"reviewer_1": "No further comments", "reviewer_2": "I co-reviewed this manuscript with one of the reviewers who provided the listed reports. This is part of the Nature Communications initiative to facilitate training in peer review and to provide appropriate recognition for Early Career Researchers who co-review manuscripts.", "reviewer_4": "This seems to be a revised manuscript and I mainly focus on reviewing the response from the authors to comments of Reviewer 3. Generally, the authors adequately addressed the comments from Reviewer 3. The main concern from Reviewer 3 is that similar individual module technologies have been proposed by peers for the recycling of lithium and other metallic elements from LIB streams. In my opinion, integration of individual units to achieve stepwise separation of multiple metals with high purity is a novelty. As the individual technologies have been proposed, an integrated train with multi-objectives and high performance of recovering several metals is valuable, as the key is to design a correct sequence that this work indeed achieves.\nTo response to the reviewer's comment, the authors also applied the proposed technology train to real spent LIB, providing more evidence to the practical feasibility of the technology. I appreciate this effort that is not commonly adopted in the literature.\nIn addition, the authors modified the structure and explained more about the mechanisms of the selectivity of the technology, which enhance the quality of this work. As the authors have provided a thorough response to the comments from reviewers (79 pages), I will not add more comments considering the timely topic of this work. Thus, I consider this work acceptable by Nature Communications."}, "author_response": "We sincerely thank the reviewer for his/her efforts and recommendation.\nReviewer: 2\nThank you very much for your review and kind recommendation.\nReviewer: 4\nWe sincerely appreciate the reviewer’s thorough evaluation of our manuscript, the valuable feedback provided, and the very kind recommendation."}]

Decision letter
30-Oct.-2025
Journal: Nature Communications
Manuscript ID: NCOMMS-25-31462A
Title: “Lithium-ion battery recycling through an integrated electro-membrane
crystallization technology”
Authors: Yan Zhao, Yangbo Qiu, Lei Xia, Xi Zhang, Shuang Zheng, Gang Lu, Jin
Shang, Raf Dewil, Seth B. Darling, Bart Van der Bruggen, Chuyang Tang
Dear Reviewers,
Thank you very much for your comments and recommend our work.
Thanks a lot.
Sincerely yours,
Yan Zhao, Bart Van der Bruggen and Chuyang Tang on behalf of all authors