**High-Pressure Oxidative Leaching and Iodide Leaching Followed by Selective Precipitation for Recovery of Base and Precious Metals from Waste Printed Circuit Boards Ash**

## **Altansukh Batnasan 1,\*, Kazutoshi Haga 1, Hsin-Hsiung Huang <sup>2</sup> and Atsushi Shibayama 1,\***


Received: 1 March 2019; Accepted: 18 March 2019; Published: 20 March 2019

**Abstract:** This paper deals with the recovery of gold from waste printed circuit boards (WPCBs) ash by high-pressure oxidative leaching (HPOL) pre-treatment and iodide leaching followed by reduction precipitation. Base metals present in WPCB ash were removed via HPOL using a diluted sulfuric acid solution at elevated temperatures. Effects of potassium iodide concentration, hydrogen peroxide concentration, sulfuric acid concentration, leaching temperature, and leaching time on gold extraction from pure gold chips with KI–H2O2–H2SO4 were investigated. The applicability of the optimized iodide leaching process for the extraction of gold from the leach residue obtained after HPOL were examined at different pulp densities ranging from 50 g/t to 200 g/t. Results show that the removal efficiency was 99% for Cu, 95.7% for Zn, 91% for Ni, 87.3% for Al, 82% for Co, and 70% for Fe under defined conditions. Under the optimal conditions, the percentage of gold extraction from the gold chips and the residue of WPCBs was 99% and 95%, respectively. About 99% of the gold was selectively precipitated from the pregnant leach solution by sequential precipitation with sodium hydroxide and L-ascorbic acid. Finally, more than 93% of gold recovery was achieved from WPCB ash by overall combined processes.

**Keywords:** waste printed circuit board; gold; iodide; iodine; ascorbic acid; leaching; precipitation

## **1. Introduction**

Hydrometallurgical processing is often used to produce metals from complex ores, concentrates, mine tailings, and secondary sources containing metals, termed as raw materials that are difficult to treat by conventional mineral processing and pyrometallurgical methods [1–3]. The leaching process is an essential step in hydrometallurgical processing and can extract metals from raw materials containing metals using a solvent. A key challenge in the leaching process is to meet selective extraction of a metal of interest from raw material avoiding some issues associated with the further processing of the pregnant leach solution [4,5].

Printed circuit boards (PCBs) are the essential building block in the majority of electric and electronic equipment (EEE) and account for 3–6% of the total constitution of EEE [6,7]. Various kinds of metals involving precious metals (Au, Ag, and Pd), base metals (Cu, Fe, Al, Ni, and Co), and toxic metals (Hg, Pb, Cr and Cd), and different types of plastic materials including non-flame retardant and flame retardant polymers are used in manufacturing of PCBs in different ratios depending on the products [8,9]. It means that waste printed circuit boards (WPCBs) generated from discarded electric and electronic equipment (EEE) can be considered as a secondary source for metals. Because the proportion of valuable metals (Au, Ag, Pd, Pt, Cu, Ni, Co, Zn, and Al) is several times higher than those in their primary ore minerals [8–11], the recovery of valuable metals from WPCBs is the real challenge to maintain their supply chains and reduce the environmental consequences of metal mining. Therefore, over the years a variety of methods such as mechanical, hydrometallurgical, bio-metallurgical, and pyrometallurgical technologies have been used to recover valuable metals from WPCBs [12–14]. Among them, highly selective separation methods, such as hydrometallurgical processes, are more suitable due to the diverse and complex characteristics of the waste [2,15]. The recovery of precious metals from WPCBs has been extensively investigated because of their high economic values compared to other metals [16].

Cyanide leaching has been employed for over a century in gold extraction industries all over the world [17,18]. Nonetheless, in gold hydrometallurgy, a great deal of research has been conducted on the replacement of cyanide with alternative reagents such as thiosulfate, thiourea, and halides due to the environmental and human health risks associated with cyanide toxicity [19–22]. Among them, an iodine–iodide solution is a highly selective and efficient leaching agent for the recovery of gold from gold ores and WPCBs [21,23–25]. Several decades ago, the first patented process by Homick et al. (1976) used an iodine–iodide solution and a reducing agent (hydrazine) to recover gold from PCBs [26]. Subsequently, numerous studies have been devoted to developing the iodine–iodide leaching process for the recovery of gold from gold chips, gold ores, and WPCBs [27–29]. Nevertheless, the main obstacles of the leaching process are high reagent consumptions, higher reagent costs, and the complex chemistry of gold leaching in iodide solution. For the last several decades, few authors have studied the dissolution of gold in iodide solution with several oxidants namely hypochlorite, oxygen, and hydrogen peroxide, respectively to reduce iodine consumption [30–32]. Although several studies have focused on the dissolution of gold in iodine–iodide and iodide solutions with different oxidants, much less attention has paid to recover gold from pregnant leach solution (PLS) resulted from leaching [33–35]. Therefore, extensive studies on the process optimization, the dissolution mechanism of gold in iodide solution, and gold recovery from PLS are necessary.

In this study, gold recovery from the WPCBs ash by a two-step leaching process involving HPOL and iodide leaching followed by reductive precipitation was investigated. The removal of base metals from the WPCBs ash by high-pressure oxidative leaching was examined at different temperatures ranging from 100 ◦C to 180 ◦C. The dissolution behavior of gold from gold chips in iodide solutions with hydrogen peroxide and sulfuric acid was studied aiming to optimize the leaching process. Various experimental parameters such as iodide concentration, hydrogen peroxide concentration, sulfuric acid concentration, leaching time, pulp density, and leaching temperature were optimized. The feasibility of extracting gold from the leach residue obtained from HPOL was evaluated at the defined KI–H2O2–H2SO4 leaching conditions under different pulp densities. The gold recovery from the pregnant leach solution via sequential precipitation using sodium hydroxide and ascorbic acid was also discussed.

## **2. Materials and Methods**

## *2.1. Materials*

Waste printed circuit boards (WPCBs) ash received from the company (Dowa Metals and Mining Co., Ltd., Akita, Japan) and commercially available pure gold chips (99.99%, 10 × <sup>10</sup> × 1 mm3 in size, AUE02CB) purchased from the chemical company (Kojunda Chemical Laboratory, Co. Ltd., Sakado, Japan) were used as starting materials in this study. The WPCBs ash was crushed and ground to pass through a 100 μm (140 mesh) sieve using a jaw crusher (Pulverisette 1, Fritsch, Fritsch GmbH & Co. KG, Idar-Oberstein, Germany) and a grinder (HERZOG, HP-M 100, Maschinenfabrik GmbH & Co. KG, Osnabrück, Germany).

The pure gold chips were first flattened and then cut into small pieces (2 × <sup>2</sup> × 1 mm3) and were used in iodide leaching to optimize the process.

The chemical composition of the WPCBs ash sample was analyzed with inductively coupled plasma optical emission spectrometry (ICP-OES, SPS-5510, Seiko Instruments Inc., Tokyo, Japan) and X-ray fluorescence (XRF, ZSX Primus II, Rigaku Corporation, Tokyo, Japan). The results obtained are presented in Table 1.

**Table 1.** The chemical composition of the waste printed circuit boards (WPCBs) ash sample.


The particle size distribution of the sample of WPCBs ash was characterized using a particle size analyzer (Microtrac, MT3300EXII, Nikkiso Group, Osaka, Japan) (Figure 1). The main components in the sample identified using X-ray diffractometer (XRD, RINT-2200/PC, Rigaku, Tokyo, Japan) were of copper oxide (CuO), quartz (SiO2), aluminum oxide (Al2O3), and tin oxide (SnO2), as shown in Figure 2.

**Figure 1.** Particle size distribution of waste printed circuit boards (WPCBs) ash.

**Figure 2.** X-ray diffraction pattern of WPCBs ash.

Sulfuric acid (H2SO4), potassium iodide (KI), iodine (I), hydrogen peroxide (H2O2), sodium hydroxide (NaOH), and L-ascorbic acid, L-AA (C6H8O6) were purchased from Tokyo Chemical Industry Co., Ltd. (Tokyo, Japan) and Wako Pure Chemical Industries, Ltd. (Osaka, Japan). All the chemical reagents were analytical grade and used as received. Distilled water was used to prepare all aqueous solutions: 1 M H2SO4,1MH2O2, 18-72 mM KI, 0.1 M NaOH, and 0.1 M L-AA. Chemical composition and constituents of the WPCB ash sample, solid residues from leaching, and precipitates were characterized using XRD, XRF, and field emission-scanning electron microscope combined with energy dispersive x-ray spectroscopy (FE/SEM-EDS), JSM-7800F (JEOL, Tokyo, Japan). Concentrations of metals in the aqueous phases from leaching and precipitation were determined using ICP-OES (Seiko Instruments Inc., Tokyo, Japan). The changes of pH and oxidation-reduction potential (ORP) of the solutions through leaching and precipitation of metals under different media were examined by pH/ORP meter (Laqua, D-74, Horiba, Ltd., Kyoto, Japan). A vacuum pump (Buchi, V-700, BUCHI Labortechnik AG, Flawil, Switzerland) and a mini centrifuge (Cat. Number C0302, Argos Technologies, Inc., Elgin, IL, USA) were used to separate solids from liquid phases after leaching and precipitation experiments.

## *2.2. Leaching Experiments*

The process optimization, separation, and extraction of the base and precious metals from the WPCBs ash via a two-step leaching procedure involving high pressure oxidative leaching (HPOL) and iodide leaching are described in this section.

## 2.2.1. Optimization of Iodide Leaching

Gold leaching experiments were carried out according to the following procedure: a certain amount of small pieces (2 × <sup>2</sup> × 1 mm3) of pure gold chips was inserted in a constant volume (10 mL) of an aqueous solution consisting of a mixture of iodide, hydrogen peroxide, and sulfuric acid (KI–H2O2–H2SO4) into a lidded 30 mL Teflon beaker, and placed in a glass vessel containing water onto a hot plate. In order to optimize the gold leaching in the system, various relevant variables, which are iodide concentration (18–72 mM KI), H2SO4 concentration (0–100 mM), H2O2 concentration (5–30 mM), leaching temperature (20–80 ◦C), and leaching time (2–12 h), were investigated under various intervals at stirring speed of 550 rpm.

After iodide leaching, undissolved pure gold species were separated from the liquor solution by filtration and dried into a drying oven at 80 ◦C for 24 h. The efficiency of gold extraction was calculated by weighing the amount of gold undissolved and comparing the weight loss of gold to the initial gold weight used in a leaching experiment. The liquor solution was analyzed by ICP-OES (Seiko Instruments Inc., Tokyo, Japan) for determination of gold concentration.

## 2.2.2. High Pressure Oxidative Leaching

Leaching experiments were carried out in an autoclave equipped with a 200 mL Teflon vessel (Nitto High Voltage Co., Ltd., Tsukuba, Japan), impeller, inlet, and outlet gas pipes, pressure gauge and temperature sensor, and electrical heating system (Figure S1). An amount of 10 g of a powder sample of WPCBs ash was firstly mixed with 100 mL, 1 M H2SO4 solution into a vessel, and placed the prepared slurry into the autoclave. Then, the autoclave system was heated up to the selected temperature (100–180 ◦C) and set to an impeller speed of 750 rpm. After that, pure oxygen gas was injected to the slurry into the vessel by adjusting the total pressure (*P*tot) of 2 MPa using a pressure gauge and held for 30 min. The total pressure (*P*tot) includes a partial pressure of oxygen gas (*P*ox) and vapor pressure (*P*v) into the vessel in the autoclave. After the leaching experiment, the autoclave was cooled down up to room temperature, and the slurry from the leaching was filtered using a membrane filter (Toyo Roshi Kaisha, Ltd., Tokyo, Japan) by a vacuum pump. The composition of leachate and solid residue generated from HPOL were analyzed by various techniques to estimate the extraction efficiency of metals.

## 2.2.3. Iodide Leaching for Extraction of Gold from WPCBs Ash

The leach residue of WPCBs ash obtained from HPOL was used as a sample to extract precious metals especially gold via iodide leaching. The leaching experiments were conducted at varying pulp densities ranging from 50 g/L to 200 g/L under the optimized conditions explained in Section 2.2.1 (72 mM KI, 20 mM H2SO4, 15 mM H2O2, stirring speed of 550 rpm, at 60 ◦C for 8 h). After completing the leaching period, solid and liquid phases were separated using the filtration with the membrane filter. The solid residue was analyzed by XRD and XRF, respectively. The concentrations of metals in a pregnant leach solution were determined by ICP-OES. The efficiencies of metals dissolution in the iodide solution and accompanying metals remained in the residue were calculated by mass balance.

## *2.3. Evaluation of Effectiveness of Iodide Leaching*

The extraction of precious and base metals from the residue of WPCBs ash via iodide (KI–H2O2–H2SO4) leaching was evaluated with the comparison of iodine–iodide (I–KI–H2O) leaching. The experiments were carried out using 20 mM H2SO4, 15 mM H2O2 in the iodide solution, and 0.08 mM I2 in the iodine–iodide solution under the same condition of KI concentration (72 mM), pulp density of 50 g/L, temperature (60 ◦C), leaching time (8 h), and stirring speed (550 rpm).

## *2.4. Precipitation of Gold from Pregnant Leach Solution*

The precipitation experiments were conducted in 20 mL of plastic tube at room temperature (25 ◦C) at a stirring speed of 550 rpm for 10 min. Firstly, minor amounts of metal impurities in the pregnant leach solution were selectively removed by precipitation under the alkaline conditions using a 0.1 M NaOH solution. Then, gold that remained in the resulting alkaline solution was reduced into its elemental form in acidic conditions using 0.1 M L-AA, which is known as vitamin C as a dietary supplement. After the precipitation experiment, a colloidal solution centrifuged at 9200 rpm for 5 min and the precipitate obtained was washed with 5 mL distilled water and then dried at 80 ◦C for 24 h. The precipitates were characterized by SEM-EDS. To determine the efficiency of metal precipitation, the supernatant solution from the centrifugation was analyzed by ICP-OES.

## **3. Results and Discussion**

## *3.1. Optimization of Iodide Leaching*

To properly and better optimize the iodide leaching process, pure gold chips dissolved in a mixture composed of potassium iodide, hydrogen peroxide, and sulfuric acid (KI–H2O2–H2SO4), preventing an interference effect of accompanying metals existing in WPCBs on gold extraction.

The generation of iodine species via oxidation of iodide with hydrogen peroxide in acidic media and the dissolution of gold in the solution are represented by the following Equations (1)–(5) [27–31,36,37]:

$$\text{H}^{\cdot}\text{I}^{-} + \text{H}\_{2}\text{O}\_{2} + 2\text{H}^{+} \rightarrow \text{I}\_{2} + 2\text{H}\_{2}\text{O} \qquad \qquad \text{E} = 1.143 \text{ V} \tag{1}$$

$$\text{I}\_2 + \text{I}^- \rightarrow \text{I}\_3^- \qquad \qquad \text{K} = 713 \tag{2}$$

$$\text{Au} + 2\text{I}^- \rightarrow \text{AuI}\_2\text{I}^- + \text{e}^- \qquad \quad \quad \quad \quad \quad \quad \quad \quad \quad -0.578 \text{ V} \tag{3}$$

$$\text{Au} + 4\text{I}\_3\text{}^- \rightarrow 3\text{AuI}\_4\text{}^- + \text{e}^- \qquad \quad \quad \quad \quad \quad \quad \quad = -0.757\text{ V} \tag{4}$$

$$2\text{Au} + \text{I}^- + \text{I}\_3^- \rightarrow 2\text{Au} \text{I}\_2^- \qquad \qquad \text{E} = -0.042 \text{ V} \tag{5}$$

Leaching parameters such as potassium iodide concentration, hydrogen peroxide concentration, sulfuric acid concentration, leaching temperature, and leaching time that affect the dissolution of gold in the KI–H2O2–H2SO4 system were investigated, and results obtained are summarized in the following sections.

## 3.1.1. Effect of Potassium Iodide Concentration and Hydrogen Peroxide Concentration

The gold extraction was examined under different concentrations of KI (18–72 mM) and H2O2 (5–30 mM) in dilute aqueous solutions of H2SO4 (40 mM). The leaching experiments were performed at 40 ◦C, 550 rpm for 2 h. The effects of KI and H2O2 concentrations on gold extraction from the gold chips were exhibited in Figure 3. Results showed that gold extraction greatly varied with variations in the concentration of KI and H2O2, respectively. At 18 mM KI concentration, gold extraction decreased greatly from 84.2 mg/L to 0.2 mg/L as H2O2 concentration increases from 5 to 30 mM. Whereas at KI concentrations of 48 mM and 72 mM, rapid increases in the gold concentration from 113.3 mg/L to 180.7 mg/L and 166.5 mg/L to 279 mg/L were observed with the addition of 5–10 mM and 5–15 mM H2O2, respectively. When the H2O2 concentration increased further up to 30 mM, the extraction of gold decreased to 34.2 mg/L and 112.8 mg/L, respectively. It appears quite clearly in Figure 3 that the maximum gold concentrations of 84.2 mg/L, 180.7 mg/L, and 279 mg/L were achieved after 2 h leaching at 40 ◦C in KI–H2O2–H2SO4 solutions composed of 18 mM KI, 40 mM H2SO4, and 5 mM H2O2; 48 mM KI, 40 mM H2SO4, and 10 mM H2O2; and 72 mM KI, 40 mM H2SO4, and 15 mM H2O2, respectively. It implies that the highest gold extraction achieved with KI–H2O2 molar ratios within the range of 4:1–5:1. As the molar ratio of KI–H2O2 was less than 3:1, the gold extraction was decreased. Results indicate that the appropriate molar amounts of KI and H2O2 in the leaching mixture is the most important to generate a suitable amount of tri-iodine species (I3 −) that promotes the gold extraction from the gold sample. It suggests that the deficient amounts of H2O2 (5–10 mM) lead to produce insufficient concentrations of I3 −; on the contrary, the excessive amounts of H2O2 (30 mM) result in the formation of solid iodine (I2(s)) under oxidizing conditions as indicated by Equation (6):

$$\text{H}^{\cdot}\text{H}^{\cdot}\text{(aq)} + \text{H}\_{2}\text{O}\_{2(aq)} + \text{H}^{\cdot}\text{(aq)} \rightarrow \text{I}\_{2(s)} + 2\text{H}\_{2}\text{O}\_{(l)}\qquad\text{(high oxidation condition)}\tag{6}$$

**Figure 3.** Gold extraction as functions of KI and H2O2 concentrations in the KI–H2O2–H2SO4 solution.

As a result, 72 mM KI and 15 mM H2O2 were selected as the suitable molar amount (5:1) between KI and H2O2 for the subsequent experiments.

## 3.1.2. Effect of Sulfuric Acid Concentration

Figure 4 shows the effect of sulfuric acid concentration on the gold extraction from pure gold chips. The concentration of H2SO4 in the leaching solution consisting of 72 mM KI and 15 mM H2O2 was varied between 0 (without H2SO4) and 100 mM. Other experimental parameters were fixed as described above.

**Figure 4.** Gold extraction as a function of H2SO4 concentration in the KI–H2O2–H2SO4 solution.

Results indicate that the absence of H2SO4 in the leaching solution results in lower gold extraction (5 mg/L), whereas the presence of H2SO4 leads to a positive effect on the gold extraction. At 20 mM H2SO4 concentration, gold extraction increased significantly to a maximum of 292 mg/L Au. Whereas an increase in H2SO4 concentration further up to 100 mM resulted in a decrease in Au concentration to 214.6 mg/L. It suggests that the efficiency of the gold extraction depends on the molar ratio of the reactants in the leaching solution. It is estimated that the molar ratios of KI–H2O2 –H2SO4 in the solutions were in the range between 5:1:0 and 5:1:7. The maximum gold extraction of 292 mg/L was achieved at the molar ratio between KI–H2O2–H2SO4 of 5:1:1. The decrease in gold extraction might be explained by the formation of a precipitate of iodine (I2(s)) due to the presence of 40–100 mM H2SO4, at which the values of pH and Eh of solutions did not change obviously, whereas these values changed drastically with 20 mM H2SO4 (Figure 5).

**Figure 5.** Effect of H2SO4 concentration on pH and Eh in the leaching solution.

As a result, the most appropriate sulfuric acid concentration for the extraction of gold was selected as 20 mM for the next experiments.

## 3.1.3. Effect of Leaching Temperature and Time

The effect of leaching temperature and leaching time on gold extraction from the gold chips were investigated under different temperature ranges from 20 ◦C to 80 ◦C and the various times from 0.25 h to 12 h. Other experimental variables such as KI concentration, H2O2 concentration, H2SO4 concentration, and stirring speed were kept constant explained in the above experiments. The results are summarized in Figure 6. It showed that the extraction of gold rises dramatically with an increase in leaching temperature and time, respectively. At the temperatures of 20 ◦C and 40 ◦C, gold extraction grew continuously for a period of leaching time (0.2–12 h), and reached its highest levels of 670 mg/L and 1127 mg/L from 18 mg/L and 30 mg/L, respectively. The maximum extraction of gold was 1137 mg/L at 60 ◦C after 8 h leaching that was equal to 99% extraction efficiency of gold. Whereas at 80 ◦C, the most effective extraction of gold had achieved 1064 mg/L for 6 h leaching that was consistent with 92.6% extraction efficiency of gold. Further increase in the temperature and time results in the decrease in the extraction of gold (Figure 6). The results revealed that the gold extraction with the KI–H2O2–H2SO4 solution is dependent on various leaching parameters such as iodide concentration, hydrogen peroxide concentration, sulfuric acid concentration, leaching temperature, and leaching time, respectively. Consequently, the best conditions for gold leaching in the KI–H2O2–H2SO4 solution determined were: KI concentration of 72 mM, H2O2 concentration of 15 mM, H2SO4 concentration of 20 mM, stirring speed of 550 rpm, leaching temperature of 60 ◦C, and leaching time of 8 h.

**Figure 6.** Gold extraction as a function of leaching time under different temperatures.

## *3.2. Extraction of Base Metals from WPCBs Ash by HPOL*

The presence of higher amounts of metal impurities in WPCBs causes adverse effects on precious metals extraction and reagent consumption [30]. To reduce the side effects of metal impurities on gold extraction and iodide consumption, HPOL experiments were carried out at different temperatures.

## Effect of Temperature

High pressure oxidative leaching as a pretreatment for recycling of base metals from the WPCBs ash was conducted at different temperatures ranging from 100 to 180 ◦C using a laboratory-scale autoclave under conditions of fixed sulfuric acid concentration (1 M), pulp density (10 g/t), agitation speed (750 rpm), total pressure (2 MPa), and leaching time (30 min). Figure 7 shows the extraction percentage of base metals from the WPCBs ash under the conditions. Results showed that the maximum extraction of Cu (99%), Zn (95.7%), Ni (91%), Al (87.3%), Co (82%), and Fe (70%), which are main base metals in the pulverized WPCBs ash, were achieved in 1 M H2SO4 at 160 ◦C, while only 1.3% of Pb was dissolved. The reaction between metal oxides and a dilute acid is often quite slow because the oxidant (H+, SO4 <sup>2</sup>−) is not enough to oxidize the metal. However, it is possible to take place the progress of the reaction in the presence of oxygen that acts as an oxidizing agent. The dissolution of base metals with dilute sulfuric acid in the presence of oxygen is, hence, represented as follows [38,39]:

$$\text{CuO} + \text{H}\_2\text{SO}\_4 \rightarrow \text{CuSO}\_4 + \text{H}\_2\text{O} \tag{7}$$

$$\text{ZnO} + \text{H}\_2\text{SO}\_4 \rightarrow \text{ZnSO}\_4 + \text{H}\_2\text{O} \tag{8}$$

$$\text{CoO} + \text{H}\_2\text{SO}\_4 \rightarrow \text{CoSO}\_4 + \text{H}\_2\text{O} \tag{9}$$


$$\text{Fe}\_2\text{O}\_3 + 3\text{H}\_2\text{SO}\_4 \rightarrow \text{Fe}\_2(\text{SO}\_4)\_3 + 3\text{H}\_2\text{O} \tag{12}$$

**Figure 7.** Extraction percentage of base metals as a function of temperature. (Conditions: 1 M H2SO4, 100–180 ◦C, 10 g/t pulp density, 750 rpm stirring speed, 2 MPa pressure for 30 min leaching).

It can be seen that precious metals (Au, Ag, and Pd) and tin (Sn) in the WPCBs ash were not dissolved in 1 M H2SO4 under the high pressure oxidation conditions due to their extremely slow leaching kinetics. It may require strong H2SO4 and longer leaching time to extract these metals.

The chemical composition of the leachate and the solid residue resulted from HPOL with 1 M H2SO4 are presented in Table 2.


**Table 2.** The chemical composition of the leachate and the solid residue of the WPCBs ash after high-pressure oxidative leaching (HPOL).

It shows that the concentration of precious metals in the leach residue was higher than that in the WPCBs ash. It is worth noting that the residue is a potential source of precious metals. On the other hand, the leachate contained the proper amount of base metals for recovery. It suggests that HPOL could be used to concentrate base metals in the leachate and precious metals in the leach residue, separately. Therefore, the leach residue obtained by HPOL was used to recover precious metals by iodide leaching in a further study.

The intensities of diffraction peaks corresponding to the copper phases in the residue became rather low compared to those from the initial WPCBs ash sample (Figure S2).

## *3.3. Extraction of Precious Metals from Leach Residue via Iodide Leaching*

## Effect of Pulp Density

The leach residue of WPCBs ash obtained from HPOL was dissolved in a mixture of KI–H2O2–H2SO4 under the selected conditions described in the previous section (Section 3.1). Figure 8 shows the percentage of metals extracted from the leach residue of WPCBs under the different pulp densities of 50 g/t, 100 g/t, and 200 g/t at the optimum conditions. The results show that an increase in the pulp density leads to a decrease in the percentage extraction of metals from the leach residue. It means that the extraction of metals had similar tendencies under the conditions.

**Figure 8.** The percentage extraction of metals as a function of pulp density.

It is observed that gold and copper are preferentially extracted from the residue via KI–H2O2–H2SO4 leaching. The percentage extraction of gold and copper decreased from 98% to 84% and from 45% to 25%, respectively, with an increase in the pulp density, whereas the percentages of both silver and palladium extraction were lower than 1%. However, the percentage extraction of metals at high pulp density reduced obviously than that in lower pulp densities, and the concentrations of metals in pregnant leach solutions increase with an increase in pulp density, as shown in Table 3.


**Table 3.** Concentration of metals in the pregnant leach solutions after leaching, mg/L.

## *3.4. Evaluation of Effectiveness of Iodide Leaching*

Figure 9 presents the results of a quantitative comparison of the effectiveness of the iodide (KI–H2O2–H2SO4) and iodine–iodide (I2–KI–H2O) leaching for precious and base metals extraction from the leach residue of WPCBs ash under the same leaching conditions.

**Figure 9.** Comparison of the efficiency of metals extraction via iodide leaching and iodine–iodide leaching at pulp density of 50 g/L.

The results show that the extraction of gold and lead by leaching with KI–I2–H2O are somewhat higher than those achieved with KI–H2O2–H2SO4. On the contrary, the efficiencies of other metals (Cu, Zn, Ni, Al, and Fe) extraction with KI–H2O2–H2SO4 were slightly higher than that in KI–I2–H2O due to the existence of SO4 − ions in KI–H2O2–H2SO4 leaching system. It was observed that no significant differences resulted in the gold extraction from the sample by both iodide and iodine–iodide leaching. It suggests that iodide leaching is as much an effective method as iodine–iodide leaching for the extraction of gold from samples containing gold and can potentially reduce reagent cost, particularly in iodine consumption.

## *3.5. Recovery of Gold from the Pregnant Leach Solution*

The pregnant leach solution (PLS) containing 32.3 mg/L Au and considerable amounts of base metal impurities such as Cu, Zn, Ni, Al, Fe, and Pb was used in the subsequent study (Table 3). The PLS has a pH of 1.9 and Eh of 0.64 V. The recovery of gold from the PLS consists of two steps: (1) removal of metal impurities via precipitation using NaOH; (2) gold recovery from the resulting solution by reductive precipitation using L-AA as a reducing agent. The results of the sequential precipitation are presented in the following sections.

## 3.5.1. Removal of Metal Impurities from the PLS

Based on the preliminary experimental results [25,40], the pH of the PLS was adjusted up to 9.0 from 1.9 using 0.1 M NaOH at ambient temperature and stirring speed of 500 rpm. The efficiency of metals precipitation is shown in Figure 10. The results show that the vast majority of Cu (99.6%), Zn (100%), Ni (90.4%), Al (100%), Fe (100%), and Pb (97%) as metal impurities were precipitated from the PLS, whereas gold did not precipitate under the condition. It indicates that most gold (32 mg/L) and trace amounts of Cu (618 μg/L), Ni (665.6 μg/L), and Pb (237.3 μg/L) have remained into the resulting solution which had a pH of 9.0 and Eh of 0.21 V. It implies that this process is efficient in removal of metal impurities from PLS.

**Figure 10.** Removal of metal impurities from the pregnant leach solution (PLS) by precipitation at pH 9.

## 3.5.2. Recovery of Gold from the Solution by Reductive Precipitation Using L-AA

Metals remained in the alkaline solution that resulted from the precipitation of base metals with NaOH were recovered by reductive precipitation using a 0.1 M, L-AA solution under conditions that were fixed as the molar ratio between Au and L-AA of 1:1, temperature of 25 ◦C, stirring speed of 500 rpm and precipitation time for 10 min [25,40]. Figure 11 shows the recovery of metals like Au, Cu, Ni, and Pb from the solution via precipitation with L-AA.

**Figure 11.** Recovery of metals from the solution by reductive precipitation with L-AA.

The result showed that gold recovery of 99.2% was achieved from the solution via the precipitation with L-AA under the conditions, while yields of Cu (0.4%), Ni (1.4%), and Pb (0.1%) as metal impurities in the precipitate were negligible. It was observed that the addition of L-AA results in a reduction of pH of the solution from 9 to 1.8 and an increase of Eh from 0.21 V to 0.64 V. The Eh-pH diagram for gold iodide and ascorbic acid species in the solution was constructed at 25 ◦C within the pH range from 0 to 14 using STABCAL software (N NBS, Helgeson thermodynamic data source, W32-Stabcal Version 1.0, Montana Technological University (Montana Tech), Butte, MT, United States) as shown in Figure 12.

**Figure 12.** The Eh–pH diagram of gold–iodide and ascorbic acid species in the solution. (Condition: [Au] = 32.3 mg/L, [I] = 0.072 M (72 mM), Eh = −0.5–1.5 volts, *P* = 1 atm at pH 0–14, 25 ◦C, (NBS, STABCAL software)).

The thermodynamic data for gold iodide and L-AA related species were imported from the references [25,41–43]. Table S1 summarizes the thermodynamic data for individual species in the solution used to construct the Eh–pH diagram. As shown in Figure 12, L-AA/H2Asc is converted into its oxidized forms such as HAsc−, Asc2−, and DHA at the different pH conditions as indicated by Equation (13) [43–45].

$$\text{H}\_2\text{Asc} \rightleftharpoons \text{HAsc}^- \rightleftharpoons \text{Asc}^{2-} \rightleftharpoons \text{DHA} \tag{13}$$

It implies that DHA-dehydroascorbic acid (C6H6O6) is the most stable oxidized product of L-AA in the pH range from 0 to 14.

These findings suggest that colloidal gold can be obtained from gold–iodide solutions by reductive precipitation with L-AA. The mechanism of the reduction of gold iodide complexes via L-AA is generally explained by the Equations (14)–(16) [25,46]:

$$\text{C}\_6\text{H}\_8\text{O}\_6 = \text{C}\_6\text{H}\_6\text{O}\_6 + 2\text{H}^+ + 2\text{e} \qquad \quad \text{E} = 0.35\text{ V} \tag{14}$$

$$\text{AuI}\_4\text{}^- + 2\text{ë} = \text{AuI}\_2\text{}^- + 2\text{I}^- \qquad \qquad \text{E} = 0.551 \text{ V} \tag{15}$$

$$\text{AuI}\_2\text{}^- + \bar{\text{e}} = \text{Au}^0 + 2\text{I}^- \qquad \text{E} = 0.578 \text{ V} \tag{16}$$

The colloidal gold as a precipitate formed from the reductive precipitation was analyzed by using FE/SEM-EDS. The SEM result shown in Figure S3 indicates that produced particles do not have a definite shape and size due to the agglomeration of particles formed. The EDS analysis in the area identified a sharp peak corresponding to gold (Figure S4).

## **4. Conclusions**

This study aimed to recover gold from WPCBs ash by hydrometallurgical processes involving HPOL and potassium iodide leaching (KI–H2O2–H2SO4) followed by sequential precipitation using NaOH and L-AA. The following conclusions can be drawn based on the results obtained from this study:


As a result, a multi-stage hydrometallurgical procedure with greater than 93% gold recovery from WPCBs is proposed.

**Supplementary Materials:** The following are available online at http://www.mdpi.com/2075-4701/9/3/363/s1, Figure S1: A schematic diagram of an autoclave used in this study. Figure S2: A comparison of the XRD pattern of the leach residue of HPOL with WPCBs ash sample. Figure S3: FE/SEM image of gold particles precipitated from the solution by reductive precipitation with L-AA. Figure S4. EDS spectra of gold particles precipitated from the solution by reductive precipitation with L-AA. Table S1: The thermodynamic data for individual species for the construction of an Eh-pH diagram.

**Author Contributions:** Conceptualization, A.B. and A.S.; Methodology, A.B.; Software, A.B. and H.-H.H.; Validation, A.B., K.H., H.-H.H., and A.S.; Formal analysis, A.B. and K.H.; Investigation, A.B. and K.H.; Resources, A.S.; Data curation, A.B. and K.H.; Writing—original draft preparation, A.B.; Writing—review and editing, A.B., K.H. and A.S.; Visualization, A.B.; Supervision, A.B., K.H., and A.S.; Project administration, A.S.; Funding acquisition, A.S.

**Funding:** This research was funded by JSPS KAKENHI, Grant Number 16H04182, Japan and The APC was funded by 16H04182.

**Acknowledgments:** This work was mainly supported by the JSPS KAKENHI Grant Number 16H04182 and partially supported by the Program for Leading Graduate Schools, New Frontier Leader Program for Rare-Metals and Resources, JSPS. The authors gratefully acknowledge their financial support.

**Conflicts of Interest:** The authors declare no conflict of interest.

## **References**


© 2019 by the authors. Licensee MDPI, Basel, Switzerland. This article is an open access article distributed under the terms and conditions of the Creative Commons Attribution (CC BY) license (http://creativecommons.org/licenses/by/4.0/).

## *Article* **Valorization of Mining Waste by Application of Innovative Thiosulphate Leaching for Gold Recovery**

## **Stefano Ubaldini 1,\*, Daniela Guglietta 1, Francesco Vegliò <sup>2</sup> and Veronica Giuliano <sup>3</sup>**


Received: 18 January 2019; Accepted: 23 February 2019; Published: 28 February 2019

**Abstract:** The metals and industrial minerals contained in the tailings of mining and quarrying activities, can degrade natural environments as well as human health. The objective of this experimental work is the application of innovative and sustainable technologies for the treatment and exploitation of mining tailings from Romania. Within this approach, the recovery of high grade raw materials to be placed on the market is achieved and reduction of these wastes volume are achieved. The current study is focused on hydrometallurgical process for the recovery of gold. The innovative treatment chosen is the thiosulphate process that, compared with the conventional cyanide, has several advantages (e.g., it is more ecologically friendly and is not toxic to humans). The conventional cyanidation process shows operating limits in the case of auriferous refractory minerals, such as Romanian wastes, the object of the study. An important characteristic of thiosulphate leaching process it has the best selectivity towards gold; it does not attack the majority of the gangue mineral constituents. Gold extraction of 75% was obtained under ambient conditions of temperature. Moreover, the overall process achieved about 65–67% Au recovery, this being in line with the conventional cyanidation process. As these results are obtained by application of the thiosulfate process on a low gold content ore, they may be considered encouraging. The optimization of process parameters and operating conditions, should permit the best results in terms of process yields to be achieved.

**Keywords:** mining waste; hydrometallurgical processes; leaching kinetic; thiosulphate leaching; electrowinning; gold

## **1. Introduction**

Several studies have shown that mining activities can significantly damage, pollute and alter the environment (e.g., soil, sediments, water and air quality) as well as human health [1–4]. Furthermore, mining involves the creation of surface features that are unstable, prone to landslides and collapses, but, especially, expose the environment to exogenous vast areas of mineralized rocks and byproducts of mineral treatment. These outcomes produce significant changes in the chemical environment [5,6]. The large volumes of wastes produced by mining occupy huge areas; these accumulations can substantially change the original landscape.

Alongside mining, mining wastes, if managed properly, can be important to the economic growth of countries. The application of a proper treatment ensures environmental sustainability and minimize the risks to human health. In fact, the deposits of mining tailings should not be treated as inert wastes, but as a neo-mine: only a short-sighted vision would be limited solely to neutralize the harmful and toxic waste stored and to put in place procedures and interventions that allow the restoration of the original environmental conditions.

Instead, the enhancement of the tailings, in addition to the sale of raw materials grade high (precious metals such as gold), allows a complete and effective recovery of environmental conditions sustainable because the recycling activity also determines an economic convenience for the treatment.

Romania, like many other nations affected by a long history of mining, is now grappling with environmental and social issues related to tailings produced by mining activities. On the Romanian territory a total of 300 tailings deposits, produced by exploitation of different minerals, have been inventoried. All are of significant proportion and should be submitted to amelioration procedures, such as the neutralization of treatment residues [7,8]. These waste materials are a result of mining activities that have reached the end of the production cycle due to depletion of the reservoir or, after stopping of the process for environmental reasons (i.e., regulations put in place at the time of the Convention for Romania's entry into the EU). In Romania, the administration of these landfills is carried out by the Romanian National Agency for Mineral Resources (ANMR). The mining sites considered within this study are Bălan, Deva deposit 1, Deva deposit 2, Brad Ribita and Brad Criscior. Mine Bălan, for example, which ceased its operations in 2006, under the Convention for the entry of Romania into the EU, is a tabular deposit tectonized in metamorphic rocks. The main mineral is chalcopyrite. The remaining reserves amounted to 500,000 tonnes at 0.8% of copper content. These reserves are substantially unavailable. The costs of re-mining and carrying out of modifications to enrich the mineral are too high compared to the value of resources and costs of extraction [7–10].

The deposits of tailings, left on the land during the mining of copper ore, contain 24.5 million tonnes of slag and more than 10 million cubic meters of materials that constitute a toxic and harmful waste. These deposits are located at a distance of few kilometers from the mining center (at the neighborhood of the Bălan town) [7,8].

The removable high-grade raw materials, from the tailings of Bălan, like those of the other mentioned sites, are:


The aim of this work is to apply innovative technologies for the treatment and exploitation of Romanian mining tailings. In particular, this study is focused on the development of an hydrometallurgical process for the recovery of gold from solid wastes of mining industry. The innovative treatment chosen was the thiosulphate process, that has advantages over the conventional cyanide and is non-toxic to humans; in fact, it is environmental impact is lower than for cyanidation [11–18].

The kinetics of the thiosulphate process has been studied with the main aim to improve the characteristics that distinguishes from the conventional cyanide process for gold recovery. The main potential advantages of this innovative treatment for gold recovery, can be summarized in the following points:


6. possibility for recovery of the dissolved gold by known techniques, such as carbon adsorption and electrodeposition.

The greatest criticality of the process is constituted by the chemistry of the ammonia-thiosulphate system, that is very complicated due to the simultaneous presence of complexing ligands such as ammonia and thiosulphate, the Cu(II)–Cu(I) redox couple and the possibility of oxidative decomposition reactions of thiosulphate involving the formation of additional sulphur compounds such as tetrathionate [11].

The ultimate purpose of this work is to develop a process scheme on the basis of results obtained at laboratory scale. This will be used to perform a preliminary study of the technical feasibility of process. Figure 1 shows an example of a mining site with tailings from Romania.

**Figure 1.** Tailings of a mining site in Romania.

## **2. Materials and Methods**

## *2.1. Sampling*

The mining sites under study were: Bălan, Deva deposit 1, Deva deposit 2, Brad Ribita and Brad Criscior. For each mining site homogeneous and representative samples were prepared for subsequent characterization. To this end, after being dried in oven at 80 ◦C for one day and sieved to 4 mm, the fractions were then homogeneously and representatively sampled using a RETSCH rotary splitter. From each site, eight samples were collected.

For the characterization of the particles sizes, the prepared samples were quartered with a manual sampler and submitted to wet sieving using sieves of following sizes: 0.5 mm, 0.351mm, 0.250 mm, and 0.125 mm. The obtained granulometric fractions were filtered, dried in a laboratory oven at 80 ◦C for one day and then weighed. On the basis of the obtained weights, the distribution curves were constructed (data not showed here).

For each mineral deposit, one of the eight prepared samples was submitted to gravimetric separation by a flow table. The four fractions obtained (light, intermediate, mixed and heavy) were filtered, dried in a stove at 80 ◦C for a day and then weighed.

This rotary splitter allows the separation of sample into various fractions whose composition corresponds exactly to that of the initial sample, because only a representative sample of the initial rate can provide significant analytical results.

This procedure ensures a high degree of accuracy and reproducibility. It is used in combination with the vibrating feeder RETSCH DR 100, utilized for the homogeneous and uniform assay during the conveying of the material; this is an automatic process of sampling, without interruptions and loss of material. The speed was monitored and kept constant.

A planetary ball mill agate mod. FRITSCH pulverisette, was used for the fine grinding of the samples. The jars and grinding balls were made of agate to prevent samples contamination.

## *2.2. Characterization*

The mineralogical characterization was carried out by the technique of X-ray diffraction (X-ray diffractometer Bruker, mod. D8 Advance).

The analytical determination of metals and gold content of the fractions obtained from the table, was carried out by Perkin Elmer, mod. 400 optical plasma spectrometer (ICP-AES) with data station. This analysis was performed on the solutions achieved after sample chemical dissolution.

Homogeneous and representative samples of approximately 10 g, submitted to chemical attack, were prepared by rotary and manual splitters and subsequently milled in a planetary mill with agate jars and balls. The milling was performed till the particles sizes was less than 80 μm, this being appropriate for the subsequent leaching tests [11]. The sample for gold recovery tests was chosen considering the gold content.

The experimental work was carried out on the heavy fraction constituted by a mix between Brad Ribita and Brad Criscior samples, which had an average gold content of 3 g/t.

## *2.3. Physical Process*

The samples from mining sites Brad Ribita and Brad Criscior, chosen for their higher gold content, were sieved to −0.5 mm, and the fractions submitted to gravimetric separation by the flow table. The table was set with the aim to obtain a richer heavy fraction The heavy fractions recovered were filtered and dried in a stove at 80 ◦C for one day.

## *2.4. Grinding*

The heavy fractions, obtained by gravimetric separation, were subjected to comminution (<80 μm), using a bar mill, to make them suitable to determination of the content of gold and for the subsequent leaching tests. The grain sizes obtained were analyzed with the SYMPATEC laser granulometer.

At the end of the comminution, the drum was emptied and the slurry after being filtered was dried in a stove at 80 ◦C. The heavy fractions, after comminution, were mixed homogeneously using the rotary splitter to prepare samples of the mix Brad Ribita-Brad Criscior, and submitted to leaching tests. The gold content of the mixture of minerals was determined after chemical dissolution, with an Atomic Absorption Spectrometer (AAS Perkin Elmer mod. 460).

## *2.5. Chemical Process*

The study of the thiosulphate process, was conducted in mechanical stirred reactors made of Pyrex glass that have a capacity of 2000 mL. Leaching experiments were carried out to study the influence of the concentration of ammonia and the concentration of thiosulphate on gold dissolution, using reagents of analytical grade and distilled water.

The leaching solutions consist of sodium thiosulfate (Na2S2O3·5H2O), used as active leaching agent, ammonia (NH4OH 30%)—for the control of pH—and copper (II) sulphate (CuSO4·5H2O)—which acts as an oxidant of gold [11,16,17].

The tests were carried out at atmospheric pressure and room temperature, while the speed of mechanical agitation was kept constant at 400 rev/min. The leaching time was 4 h, the weight of the samples of 500 g, the particle size minus 80 μm, pH of 10.5 and a redox potential +0.1 V. At set time intervals, small volumes (10 mL) of leaching solution were taken from the reactor. These were analyzed to determine their Au content and, therefore for the kinetic study of gold dissolution [19,20].

The pH and the oxidation-reduction potential of the slurry, were measured using a combined glass electrode and a platinum combination electrode, respectively, both being connected to a digital pH meter.

At the end of each test, the reactor was emptied, while the filtration of the slurry was realized through pressure filters. The solid residue was submitted to washing with distilled water and ammonia; moreover, gold content was determined, after chemical dissolution of homogeneous and representative samples, with an Atomic Absorption Spectrometer (AAS Perkin Elmer mod. 460).

After leaching, the gold was purified by selective adsorption onto granular activated coconut carbon. The influence of the carbon concentration was studied.

This was conducted in Pyrex glass reactors of capacity of 2000 mL under mechanical stirring (400 rev/min), at room temperature, for a total contact time of 1 hour [20–22].

The tests were carried out by placing in contact leached solution (500 mL) with the different amounts of coconut charcoal, an activated carbon material. The influence of the contact time was investigated by performing liquor withdrawals at set intervals. The concentration of the activated carbon was varied from 5 g/L to 15 g/L. After each experiment, carbon was recovered from the solution and left air drying. Representative samples of carbon were collected and submitted to quantitative chemical analysis.

The desorption of gold from carbon was carried out by elution with a water-ethyl alcohol solution prepared using absolute ethanol (C2H5OH) [22]. The gold stripping tests were conducted in a Pyrex glass reactor, with a capacity of 250 mL. The reactor was fitted with three necks: the first one had reflux condenser for the removal of vapors, the second one housed a probe that was connected to a shaking-heating plate for stabilizing the temperature and a thermometer for temperature control was inserted through the third port. The tests were conducted varying the temperature from 40 to 85 ◦C.

Sampling was performed also at determined time intervals and then chemically analyzed to determine their Au content. These data were used to determine the kinetics of the process.

The final recovery of purified metallic gold from the water-alcohol solution, was carried out by the electrochemical process [19,20,23].

Gold recovery conducted in an electrolytic cell with a capacity of 200 ml in a jacketed Pyrex glass, connected to a Julabo, mod. 5B thermostat (control from −20 to +100 ◦C). The cell was fitted with a saturated calomel reference electrode, a working electrode (cathode) constituted by a net of platinum wire, having a surface area of 100 cm2, and a counter-electrode (anode) consisting of a spiral platinum.

The cell was connected to a AMEL, model 555 B potentiostat-galvanostat. The current flowing through the cell was converted into a numerical value by an AMEL, model 721 integrator. The potential difference between the cathode and the anode was measured with an AMEL, model 631 differential electrometer.

The presence of a magnetic stirrer bar allows stirring of the solution into the cell. The electrolysis tests were conducted using 200 mL of strip solution.

## **3. Results**

## *3.1. Physical Process*

The homogeneous and representative samples, achieved after screening, to retain particles with a diameter greater than 0.5 mm, were submitted to gravimetric separation by the flow table. The goal of the physical process was to concentrate pyrite, and then the gold associated with it, in the heavy fraction. The light fraction consists predominantly of quartz.

The heavy fraction for none of the deposits reaches 10%, therefore we must consider that most of the pyrite could be concentrated in the mixed fraction. For all sites the light fraction exceeds 40%, with over 61% in Balan samples.

The chemical analysis, allowed the gold content of the mixed fractions and heavy minerals to be studied. Based on this data the mining sites of greatest interest from the point of view of the gold content are: Brad Ribita and Brad Criscior. These sites were chosen to the study the process of Au recovery. The X-ray diffraction performed on the heavy and mixed fractions from Brad Ribita and Brad Criscior allowed the mineralogical composition to be determined. In detail, the mineralogical species contained in the heavy fraction Brad Ribita are the following: quartz—(SiO2) (48.3%); pyrite (FeS2) (20.4%); muscovite—(K, Ba, Na)0.75(Al, Mg, Cr, V)2(Si, Al, V)4O10(OH, O)2 (15.5%); albite—(Na0.75Ca0.25)(Al1.26Si2.74O8) (8.7%); chamosite—(Mg5.036Fe4.964)Al2.724(Si5.70Al2.30O20)(OH)16 (4.9%); Calcite—CaCO3 (2.1%) and Chalcopyrite—CuFeS2 (0.1%).

The following main elements have been determined by chemical analysis of the heavy fraction Brad Ribita: Si (27.0%), S (11.0%), Fe (10.6%), Al (3.3%), V (3.1%), Ca (1.2%), Ba (1.2%), Mg (0.9%), Na (0.8%).

Regarding heavy Brad Criscior, the mineralogical composition is constituted by: quartz—(SiO2) (66.2%); pyrite (FeS2) (10.2%); muscovite—(K, Ba, Na)0.75(Al, Mg, Cr, V)2(Si, Al, V)4O10(OH, O)2 (14.2%); albite—(Na0.75Ca0.25)(Al1.26Si2.74O8) (5.1%); chamosite—(Mg5.036Fe4.964)Al2.724(Si5.70Al 2.30O20)(OH)16 (2.9%); Calcite—CaCO3 (1.1%) and Chalcopyrite—CuFeS2 (0.3%).

Main elements detected by chemical analysis of the heavy fraction Brad Criscior are: Si (33.9%), S (5.5%), Fe (5.4%), V (2.9%), Al (2.5%), Ba (1.1%), Cr (0.8%), Ca (0.6%), Mg (0.7%), Na (0.5%), K (0.3%), Cu (0.1%).

FeS2 in the mixed fraction for Brad Ribita is 6.3%.

The results of the gravimetric separation are shown in the following Table 1.

**Table 1.** Heavy fractions separated from the flow table.


## *3.2. Chemical Process*

## 3.2.1. Gold Leaching

The experimental work was carried out on the heavy fraction, mix of the Brad Criscior and Brad Ribita samples, with an average gold content of 3 g/t. Gold was recovered by leaching with sodium thiosulphate as previously described.

After the preliminary experiment—carried out using conditions from the literature [11,15] leaching enabled a gold recovery of 38.10% (for the calculation of efficiency there was considered also the gold content within the washing solution and solid residue). The investigation permits the influence and the interactions between S2O3 <sup>2</sup>−, CuSO4 and NH3 concentration on gold kinetics to be studied (Figures 2 and 3).

**Figure 2.** The influence of NH3 concentration on gold recovery (2 M Na2S2O3; 0.1 M CuSO4).

**Figure 3.** The influence of S2O3 <sup>2</sup><sup>−</sup> concentration on gold recovery (4 M NH3; 0.1 M CuSO4).

Figure 2 shows the influence of NH3 concentration on gold recovery, while Figure 3 shows the influence of S2O3 <sup>2</sup><sup>−</sup> concentration. In both cases, CuSO4 concentration was constant at 0.1 M [11,14,22]: the thermodynamic condition that allows the best gold recovery was achieved with a solution of the following composition: 2 M S2O3 <sup>2</sup>−, 0.1 M CuSO4 and 4 M NH3.

Table 2 describes the best kinetic with higher gold recoveries (about 75% Au after 15 min, without washing). It was observed that the kinetics of extraction decreases by approximately 20% at the end of the experiment (about 65% Au after 4 h, with washing) [11].


**Table 2.** Gold recovery obtained with a ammoniacal leaching solution having the following composition: 2MS2O3 <sup>2</sup>−, 0.1 M CuSO4 and 4 M NH3.

The following Table 3 shows the parameters that allowed the best gold leaching kinetics.


**Table 3.** Best process parameters of the gold leaching.

The goals of the further work will be to optimize the leaching process conditions, with the aim to increase gold extraction yields, and to reduce the reagents consumption, such as the concentration of thiosulphate used.

## 3.2.2. Gold Adsorption

The leaching solution containing gold in the form of soluble complex, is placed in contact with the activated carbon to selectively separate the gold by adsorption. The experimental conditions described in Section 2, were applied to study the influence of the mass ratio of carbon to solution on the gold recovery process. The purification process of the solutions achieved after leaching allowed high gold recoveries to be obtained (Table 4).


**Table 4.** Kinetics of Au adsorption at different concentrations of carbon in solution.

The results show an almost complete recovery of the gold present in solution. From the trend it is clear that the increase of the concentration of carbon in solution favors the recovery. In particular, after 1 h, at a concentration of carbon of 5 g/L, about 86% Au was adsorbed, but the recoveries reach 99% Au when the concentration of the adsorbent increases to 10 g/L. Also it was found with a concentration of 15 g/L, that 99% Au adsorption was achieved after only 30 min.

The parameters that allowed the best gold adsorption onto activated carbon have been reported in Table 5.


**Table 5.** Best process parameters of gold adsorption on activated carbon.

## 3.2.3. Gold Desorption

The purpose of desorption was to re-extract the gold adsorbed and concentrate it. The duration of process was fixed at 6 h.

From the experimental results, shown in Table 6, it can be observed that the final gold recovery was of 79.0%: further studies using other types of alcohol such as isopropyl alcohol and ethylene glycol, may improve the efficiency, shortening the duration of the stripping process.

**Table 6.** Kinetics of Au desorption at different concentrations of carbon in solution.


The process parameters for the desorption of gold from activated carbon that permitted to achieve 99.0% Au recovery—including washing—after 6 h, are shown in Table 7.


**Table 7.** Best process parameters of gold desorption from the activated carbon.

### 3.2.4. Electrowinning

The last step of the process was electrowinning. The goal of this step was the recovery of gold from the strip solution by cathodic deposition of the metal. Table 8 shows the experimental results obtained. Table 9 shows the optimized process parameters of the phase of gold electrodeposition.



**Table 9.** Best process parameters of the phase of gold electrodeposition.


The kinetics of electrodeposition in cell of laboratory is quick and the final metallic gold recovery was high (98% Au). The concentration of gold in the sterile solution was below the detection limit of the instrument) after the first 30 min. A dark deposit was uniformly distributed on the surface of the cathode.

From the reported data, it was noted that the intensity of the measured current is around 210 mA, which corresponds to a current density of about 2.1 mA/cm2, given that the cathode surface is 100 cm2. As can be seen from the experimental results, the amount of charge that passes through the cell after 30 min was 254 Coulombs. Hence the current efficiency was low (5%); this is due to parasitic reactions, such as the reduction of the water and of dissolved oxygen. The consumption of energy was high; about 20 kWh/kg of gold deposited.

## **4. Discussion**

The preliminary study of the parameters and operating conditions for the various stages of the thiosulphate process shows, for the sample composed of mining wastes from Brad Ribita and Brad Criscior, the technical viability of the process. The experimental results obtained indicate good gold dissolution kinetics in the aqueous ammoniacal solution of thiosulphate, which can be used without special precautions and restrictions.

The gold extraction reached an average final value of 75% Au, working at room temperature, but, the trends for dissolution demonstrate that the thermodynamic parameters were not optimized, because gold recovery decreases during the best experiment. This fact, probably, is due to the

#### *Metals* **2019**, *9*, 274

thermodynamic instability of the complexing ligands (ammonia and thiosulphate) and oxidizing agents (copper ions) present in the system, tested for the first time on this type of material [11,24].

It is noted that the extraction of gold is already very high in the first 15 minutes: this is due to the gravimetric enrichment and means that a part of the gold is free and is not incorporated into the mineral matrix (see Table 2).

Gold recoveries for the overall process including leaching (extraction yield of about 75% Au) and complete adsorption-desorption–electrodeposition cycle (about 90% Au recovered) were about 65–67% (considering the gold content in the washing) in line with the conventional cyanidation process, as demonstrated in previous experimental work [25].

These results are very encouraging, considering that it is commercially an innovative process, applied to a gold ore with a low content.

Samples were leached after comminution <80 μm: a finer grinding would probably result in an increase in gold extraction, but its convenience can only be determined after careful cost analysis.

The optimization of the process is still required to identify the best process parameters and operating conditions. Considering the progressive depletion of gold mining reserves and the inability of the gold production to quickly react to the prospect of a change in prices and to changes in demand, it is providential to recover gold from mining tailings. The high price of the precious metal allows, as preliminary economic analysis shows, the feasibility of alternative processes despite the low levels of gold, the large amount of sterile to be treated and the high costs of extraction.

The enhancement of the tailings, in addition to the sale of raw materials of high grade, allows to implement an effective and sustainable recovery technology. In this way it is possible to guarantee over time the use of two indispensable resources of primary importance: the environment as a whole on one side and raw materials mining the other.

In addition to the economic and environmental aspects, we must consider the social benefits that the application of these innovative processes may generate such as providing many jobs and contributing thus to the development of repressed areas, improving the competitiveness of and creating added value and new jobs in raw materials processing, refining, equipment manufacturing and downstream industries.

In the next step of investigation, pre-treatment of components such as pyrite will be investigated before the extraction of gold. For this purpose a treatment circuit that employs biotechnological processes, with low energy consumption, will be integrated.

The elimination of pyrite will help to reduce the cost of extracting gold, reducing the consumption of the reagents. The application of physical-chemical and biological-chemical methods will allow the treatment of the Acid Mine Drainage, together with the recovery of heavy metals such as copper [26]; moreover, the optimization of leaching kinetics will be performed. The complete leaching process analysis will be outlined, including detailed description of the process scheme together with the economic analysis.

## **5. Conclusions**

The present work examined a potential innovative application of a treatment for gold recovery, based on the thiosulphate process.

The challenge here was to apply it to the recovery from resources with low gold content. The preliminary application of this circuit allowed the following results in the various steps to be obtained:


These results are very encouraging, considering that this is a commercially innovative process, applied to a low gold content ore.

The next objective is to study parameters that allow the improvement of the gold dissolution kinetics and the subsequent steps of recovery from purified solutions, thereby determining beforehand the technical feasibility of the scheme of process developed at the laboratory scale.

The optimization of process parameters and operating conditions, and scale up of the process at industrial level will permit the best results in terms of process yields to be achieved, and in turn will allow us to exploit important resources for the European economy.

**Author Contributions:** Conceptualization, F.V. and V.G.; Data curation, F.V. and D.G.; Formal analysis, S.U. and F.V.; Investigation, D.G.; Methodology, D.G.; Supervision, S.U.; Writing—original draft, S.U.; Writing—review & editing, S.U. and V.G.

**Funding:** The work was supported financially by the Istituto di Geologia Ambientale e Geoingegneria, CNR, through special funds for free theme research.

**Acknowledgments:** This experimental work was carried out during the course of the experimental degree thesis of Dr. Alessia Panone from Università degli Studi dell'Aquila. The authors are grateful to Dr. Alessia Panone and Mr. Pietro Fornari for their helpful collaboration during the experimental work.

**Conflicts of Interest:** The authors declare no conflict of interest.

## **References**


© 2019 by the authors. Licensee MDPI, Basel, Switzerland. This article is an open access article distributed under the terms and conditions of the Creative Commons Attribution (CC BY) license (http://creativecommons.org/licenses/by/4.0/).

## *Article* **Leaching Chalcopyrite Concentrate with Oxygen and Sulfuric Acid Using a Low-Pressure Reactor**

## **Josué Cháidez 1,\*, José Parga 2, Jesús Valenzuela 3, Raúl Carrillo <sup>4</sup> and Isaías Almaguer <sup>5</sup>**


Received: 7 January 2019; Accepted: 1 February 2019; Published: 6 February 2019

**Abstract:** This article presents a copper leaching process from chalcopyrite concentrates using a low-pressure reactor. The experiments were carried out in a 30 L batch reactor at an oxygen pressure of 1 kg/cm<sup>2</sup> and solid concentration of 100 g/L. The temperature, particle size and initial acid concentration were varied based on a Taguchi L9 experimental design. The initial and final samples of the study were characterized by chemical analysis, X-ray diffraction and particle size distribution. The mass balance showed that 98% of copper was extracted from the chalcopyrite concentrate in 3 h under the following experimental conditions: 130 g/L of initial sulfuric acid concentration, temperature of 100 ◦C, oxygen pressure of 1 kg/cm2, solid concentration of 100 g/L and particle size of −105 + 75 μm. The ANOVA demonstrated that temperature had the greatest influence on copper extraction. The activation energy was 61.93 kJ/mol. The best fit to a linear correlation was the chemical reaction equation that controls the kinetics for the leaching copper from chalcopyrite. The images obtained by SEM showed evidence of shrinking in the core model with the formation of a porous elemental sulfur product layer.

**Keywords:** Hydrometallurgical processes; Chalcopyrite; kinetics; low-pressure leaching

## **1. Introduction**

Chalcopyrite is the most abundant sulfide copper mineral in the Earth's crust. Generally, it is associated with other compounds such as galena, sphalerite, pyrite, arsenic, antimony or bismuth sulfides; moreover, it is often bonded with valuable metals such as silver and gold.

From an environmental and economic perspective, further technological developments for obtaining high-grade copper in an efficient and cost-effective manner are desirable. Today, companies such as Beijing Nonferrous Metal, JX Nippon Mining & Metals, Freeport McMoran, Freeport Minerals, Phelps Dodge, Outotec, BHP Billiton, etc. are investing in hydrometallurgical research because of the potential associated economic benefits [1].

Specifically, hydrometallurgical processes have a series of advantages in comparison to pyrometallurgical processes, for example, the required plant capacity is smaller (<10,000 t/y of copper), there is no need for an acid plant, no dust is emitted, etc.

Hydrometallurgical pilot plant projects such as Outotec, Galvanox, Activox and AAC/UBC have implemented some of the latest technology for mineral leaching, and several demo plants have also been installed. In particular, leaching reactors have been developed by the hydrometallurgical industry, wherein an oxidant catalyzer is commonly introduced into a pressure reactor to leach copper under varying temperatures and pressures. Table 1 lists the existing hydrometallurgical processes for leaching copper in a sulfate media, which are classified according to low, medium and high temperatures and pressures [1]. However, other aqueous media have been studied for the chalcopyrite leaching (glycine [2], nitrate [3], chloride [4], ammonium [5], etc . . . ).

Most commercial plants operate under conditions of high temperature and pressure in a sulfate medium. Nevertheless, in Las Cruces (Spain), a commercial plant with an atmospheric leaching copper process has been implemented with Outotec technology [6].

The present article focuses on the leaching stage of the hydrometallurgical process to recovery copper and iron in the liquid phase using a batch reactor under low temperature and oxygen pressure conditions. In subsequent processing, lead, silver and gold may be recovered in the resulting residue by the pyrometallurgical process of lead [7–9]. The design of this technology was based on the fundaments of thermodynamics and metallurgy.



## *Metals* **2019**, *9*, 189

## **2. Materials and Methods**

## *2.1. Material and Equipment*

The experiments were carried out in a stainless steel 316 L closed reactor having a volume of 30 liters and being equipped with an agitation system with 4 baffles, a security valve calibrated at 2 kg/cm2, a rupture disc calibrated at 3 kg/cm2, a pressure transmitter with a chemical seal and a resistance temperature detector (RTD) connected to a data logger. Also, the reactor had a controlled cooling-heating jacket. Figure 1 shows an image of the reactor.

**Figure 1.** Batch reactor for leaching chalcopyrite concentrate.

The chalcopyrite concentrate was supplied by Peñoles (Mexico). Samples of the concentrate were characterized by X-ray diffraction (XRD, Panalytical, Empyrean model), chemical analysis (CA, PerkinElmer 8300, LECO SC230DR) and a backscattered electron (BSE) module in a scanning electron microscope (SEM, FEI, Quanta600 model) for a wider range of the mineralogical species. The chemical analysis is presented in Table 2. The carbonate content was calculated with the difference of total and organic carbon. Table 3 shows the mineralogical reconstruction via XRD and CA expressed in terms of weight percentage (Wt.%). The mineralogy species obtained by BSE-SEM are shown in Table 4 in terms of weight percentage (Wt.%).




**Table 3.** Mineralogical reconstruction of the chalcopyrite concentrate.

Then, the chalcopyrite concentrate was fractionated to different sizes in a Tyler RO-TAP® Sieve Shaker using −74, −105 + 74 and −149 + 105 μm. All fractions were characterized, and no significant variation was observed in the mineralogical composition and chemical analysis. The particle size

Pyrite FeS2 7.8

distribution of the residues was measured in a Horiba LA 950 V2, which expressed the results in terms of the equivalent spherical diameter.


**Table 4.** Results of the analysis of the chalcopyrite concentrate by the SEM-BSE system.

## *2.2. Experimental Method*

To clarify the effects of particle size, temperature and initial sulfuric acid concentration on copper extraction, a Taguchi 33 experimental design was employed. In addition, a tenth test was done using different temperatures to calculate the activation energy. The following experimental conditions were constant: residence time (7 h), oxygen pressure (1 kg/cm2), solid concentration (100 g/L) and agitation velocity (550 RPM). Table 5 shows the experimental design.



The experimental procedure began with the addition of hot water (80 ◦C) to the reactor and the initiation of the agitation system, which was set at a low revolution speed; then, the chalcopyrite concentrate was fed into the reactor, followed by sulfuric acid. After the addition of these materials, a 2 min air purge was performed; then, the reactor was closed and pressurized to 1 kg/cm<sup>2</sup> with

medicinal oxygen. The agitation velocity was set at 550 RPM, and the data logger was then turned on to start recording data.

To determine the kinetics of the copper leaching process, 100 mL samples were taken from a lateral valve of the reactor at different time intervals during the test.

## **3. Results**

## *3.1. Thermodynamics*

The thermodynamics of copper sulfide leaching are based on the interaction of the elements required to carry out the decomposition of chalcopyrite. The main reactions that govern the leaching of copper concentrates are shown as follows:

CuFeS2 + 2.5O2 + H2SO4 = CuSO4 + FeSO4 + H2O+S◦ (1)

$$\text{CuFeS}\_2 + 2\text{Fe}\_2\text{(SO}\_4\text{)}\_3 = \text{CuSO}\_4 + 5\text{FeSO}\_4 + 2\text{S}^\circ \tag{2}$$

$$2\text{ FeSO}\_4 + \text{O}\_2 + 2\text{H}\_2\text{SO}\_4 = 2\text{Fe}\_2(\text{SO}\_4)\_3 + 2\text{H}\_2\text{O} \tag{3}$$

The Pourbaix diagrams in Figure 2 show that a low pH is required to keep copper and iron in sulfate solution. In addition, to ensure that ferric ions are present in the solution, the oxide potential must be above 0.57 V (SHE). This step enables indirect leaching via reaction 2, wherein the oxidation-reduction cycle of iron facilitates the decomposition of chalcopyrite. The Pourbaix diagrams were calculated with the software HSC 8.0.6 at 95 ◦C, [Cu] = 0.787 mol/L, [Fe] = 0.895 mol/L and [S] = 1 mol/L [10].

**Figure 2.** Cu-Fe-S Pourbaix diagrams.

## *3.2. Extraction and Chemical Analysis*

The results of the experiment show that the main influential variables in the leaching process were temperature and initial sulfuric acid concentration. Particle size did not significantly affect the process. The results of copper extraction versus time are presented in Figure 3.

The test number 5 presented the best results according to the mass balance. Under the corresponding conditions, 97.99% of copper was extracted in 3 h of reaction. The solid shrink was 61.8% (Wt.%), and the oxygen consumption was 0.662 g O2/g Cu fed. The density of the final solution was 1.15 g/mL, and a final oxidation-reduction potential (ORP) of 0.483V was measured in the suspension with a calomel electrode (Hg/Hg2Cl2).

Notably, different authors have reported percentages of copper extraction from chalcopyrite of 70% [11], 65% [12], 60% [13], 83% [14] and up to 95% [15]; in addition to 95% via the arbiter process, 98% via the Freeport McMoran method, 97–98% via the Activox process, 98% via the Albion process and 95% via the Galvanox process [1].

**Figure 3.** Copper extraction versus leaching time.

The test 5 was carried out at 100 ◦C with an initial sulfuric acid concentration of 130 g/L; this concentration of sulfuric acid was sufficient for reaction with species and to keep the iron in solution. Table 6 presents the CA of the residues and the solution along with the percentages of elemental distribution in the liquid phase that were obtained from the mass balance.

**Table 6.** Chemical analysis of the test 5 residue (Wt.%), chemical analysis of the solution (g/L) and the elemental distribution in the liquid phase (%) of the leaching process.


The leaching solution contained a high percentage of zinc, copper and iron because of the solubility of these elements in the utilized sulfate medium (given the temperature and acid concentration). In a global process view, it is important to consider that the solution could be treated in a solvent extraction and electrowinning stages for the production of electrolytic copper. The raffinate from solvent extraction could be neutralized with calcium carbonate to precipitate ferric ions, zinc, arsenic and minor elements; then, the solution could be recycled to the direct leaching stage.

The iron in the residue mainly corresponded with pyrite, which requires higher temperature and pressure for decomposition. Table 7 shows the species present in the solid phase in terms of weight percentages, including elemental sulfur (64.1%), anglesite (19.3%), silica (5%), pyrite (5.7%), unreacted chalcopyrite (3.23%) and gypsum (2.3%).

**Table 7.** Mineralogical reconstruction of the leaching residue in test 5.


From an economic perspective, the recovery of valuable minerals in residues following the hydrometallurgical treatment of chalcopyrite concentrates is important. The high content of elemental sulfur could reduce the profitability of recovering valuable metals by cyanidation or melting processes.

Table 8 presents the final particle size distribution of the P5 test. As expected, the particle size decreased considerably from 100% +74 μm to 90% −16.92 μm because of the leaching of chalcopyrite particles and the formation of elemental sulfur.

**Table 8.** Particle size distribution of the resulting residue from chalcopyrite leaching in test P5.


Figure 4 shows electron images obtained by SEM-BSE of unreacted chalcopyrite and galena particles, which were identified by SEM-EDS. The porous layer of elemental sulfur surrounding the particles can be observed. These particles are indicative of the shrinking core model, wherein a layer of elemental sulfur is formed as a product. However, the high extraction percentage of copper obtained in a short time (3 h) indicates that the elemental sulfur layer does not passivate the leaching of the chalcopyrite concentrate.

**Figure 4.** Punctual microanalysis by SEM-BSE-EDS in the leached residue of test P5. (**a**) Unreacted chalcopyrite particle, (**b**) unreacted galena particle, (**c**) unreacted Chalcopyrite particle, (**d**) unreacted Chalcopyrite particle.

Figure 5 shows the microstructure and the X-ray mapping by SEM-EDS for sulfur, copper and iron of the unreacted chalcopyrite particle.

**Figure 5.** X-ray mapping by SEM-EDS in the unreacted chalcopyrite particle of test P5. (**a**) electron image, (**b**) S, (**c**) Cu, (**d**) Fe.

In Figure 6, the temperature profile, in addition to the partial and accumulated oxygen consumption, are shown. At 2 h, oxygen consumption reaches its maximum and then decreases after this point.

**Figure 6.** Temperature and partial and accumulative oxygen consumption over time in test P5.

A similar finding was observed for the iron concentration in solution; ferrous ions reached their maximum concentration at 2 h and then decreased. This could explain the chalcopyrite leaching as a two steps process.

Specifically, the first step occurred within the initial 2 h of the test, wherein most of the chalcopyrite was decomposed at temperatures above 95 ◦C. The second step occurred when the remaining ferrous ions were oxidized to ferric iron. Depending on the next stages, which are related with the liquid phase (iron purification or solvent extraction), the Fe3+/Fe2+ relation must be as high as possible to ensure that the process is not affected. In Figure 7, the concentration profile of iron and ferrous ions during leaching is shown.

**Figure 7.** Concentration of iron and ferrous ions over time in test P5.

## *3.3. Statistical Analysis*

As mentioned in Table 5, this study is based on a modified Taguchi L9 experimental design with three levels for three independent variables or parameters. An additional experiment was realized (Test 10): analysis of variance (ANOVA) of the experimental tests data at different conditions was used to evaluate the effect of each individual variable. The results of Test 10 were not included in ANOVA due the difference of temperature between the lowest and the average and the small amount of Cu extraction observed.

Table 9 shows the effects of each parameter using the ANOVA module of the Minitab 15 software. The table shows the values of degree freedom (DF), sum of squares (SS), media of squares (MS), Fisher ratio (F), probability level (Prob Level) and the probability that a false null hypothesis can be rejected (Power) with a 95% confidence level (α = 0.05). According to F, Prob Level and Power values, ANOVA shows that under the studied conditions, temperature is the most important factor for copper extraction and oxygen consumption. The results also indicated that within the analyzed range, the other two variables studied (initial acid and particle size) did not have a statistically-significant effect.


**Table 9.** Analysis of variance (ANOVA) for Cu extraction.

\* α = 0.05.

Figure 8 shows the correlations of temperature, initial acid concentration and particle size with percentage of copper extraction. As observed, temperature had the greatest effect on the process of leaching copper from chalcopyrite concentrates. Figure 9 shows the correlations of temperature, initial acid concentration and particle size with oxygen consumption. As expected, temperature

once again had the greatest effect on oxygen consumption, whereas particle size and initial acid concentration had no clear influence.

**Figure 8.** Correlations of temperature, initial acid concentration and particle size with extraction of copper.

**Figure 9.** Correlations of temperature, initial acid concentration and particle size with oxygen consumption.

According to the above results the copper extraction can be calculated with the following multiple regression equation:

$$\text{Copper extraction (\%)} = -146.382 + 2.253 \text{ T} + 0.0957 \text{ Acid} - 0.0554 \text{ Size} \tag{4}$$

where T is the temperature expressed in ◦C; Acid is the initial acid concentration in g/L and size is the particle size of the chalcopyrite concentrate in μm.

## *3.4. Effect of Temperature*

As observed in Figure 3, the different tests can be categorized into three groups with differing rates of reaction that were principally determined by temperature. The tests of the first group were carried out at 100 ◦C (tests 3, 5 and 7) and resulted in 97% copper extraction within 3 h. The second group (tests 2, 4 and 9) resulted in 55–77% copper extraction within 7 h. The tests of the final group were carried out at 80 ◦C (tests 1, 6 and 8) and resulted in 10–20% copper extraction within 7 h.

To highlight the required temperature for activating the decomposition of chalcopyrite concentrates, test 5 was replicated with a slowly-increasing temperature. Figure 10 shows the percentage of copper extraction and temperature versus time. As observed, the temperature had to reach 92–95 ◦C to decompose the chalcopyrite in the concentrate.

## *3.5. Effect of Particle Size*

In the three groups that formed with respect to different reaction temperatures (Figure 3), the reactions were also faster for concentrates of small particle size; nevertheless, in the group that reacted at 100 ◦C, the −149 + 105 μm chalcopyrite concentrate reacted more rapidly than the concentrate filtered by −75 μm. The reactions in this group could have been slowed by the heat transfer from the jacket of the reactor to the suspension, resulting in different rates of reaction.

Thus, even when the chalcopyrite concentrate was a larger particle size (−149 + 105 μm), the copper extraction was not affected at 100 ◦C, and a similar level of extraction of copper was obtained.

**Figure 10.** Copper extraction and temperature for the test 5 replicated.

## *3.6. Effect of Acidity*

The initial sulfuric acid concentration in the suspension must be calculated based on the stoichiometry of the compounds that consume acid and the final acid concentration required to keep iron and copper in solution.

According to Figure 8, the initial sulfuric acid concentration is not significant for copper extraction. But in other exploratory tests carried out with the same copper concentrates at 100 ◦C under the same conditions, the suspension was found to require at least 15 g/L of sulfuric acid in solution to avoid the precipitation of iron as plumbojarosite in the residue. Thus, to prevent any problems in the recovery of valuable metals resulting from the presence of elemental sulfur and plumbojarosite in residues, and to avoid any potential impact on the profitability of operations, the initial sulfuric acid concentration is important to consider.

$$\text{2Fe2(SO4)3} + \text{12H2O} + \text{PbSO4} = \text{2Pb}\_{0.5}\text{Fe3(SO4)2(OH)6} + \text{6H2SO4} \tag{5}$$

Figure 11 shows the iron and acid concentrations in solution, the percentage of plumbojarosite in the residue and the percentage of copper extraction versus time in an exploratory test. The initial acid concentration was 72 g/L, yet it diminished to 10–15 g/L. At 15 g/L of sulfuric acid in solution, the precipitation of plumbojarosite in the residue began, leading to a clear decrease in the iron in solution from 23.6 g/L to 15.8 g/L.

**Figure 11.** Results of an exploratory test with a low acid concentration in the reaction solution.

A comparison of the copper extractions with low and high acid concentrations in the reaction solution shows that passivation was promoted by a lack of acid in the solution, which, in turn, produced a plumbojarosite layer on the chalcopyrite surface (see Figure 12).

**Figure 12.** Precipitation of plumbojarosite under low acidity conditions during leaching. (**a**) Plumbojarosite particle, (**b**) unreacted chalcopyrite particle.

The mass balance demonstrated that the ratio of sulfuric acid consumed (real) to that calculated by stoichiometry is given by Equation (6).

$$1.25 = \frac{\text{g}\,\text{Real}}{\text{g}\,\text{Stoichiometric} + \text{g}\,\text{toreach}\,\left[\text{finalacid}\right] = 45\,\text{g/L}}\tag{6}$$

## *3.7. Kinetics*

To determine the kinetics of the leaching process described herein, the shrinking core (product layer) model was applied to the real batch process considering the scanning electron microscopy (SEM) images of partly-reacted particles (Figure 4). The controlling step of the reaction was based on a comparison of the experimental data and assessment of which controlling model gives the best fit to the data. If the chemical reaction mechanism is assumed to be the controlling step, <sup>1</sup> − (1 − X)1/3 (X = conversion) is plotted as a function of time for the experimental data, and if the plot gives a linear correlation, the assumption is considered to be correct. Analogously, 1 − 3(1 − X)2/3 + 2(1 − X) (diffusion as controlling mechanism) can be plotted for the data when a non-porous product layer is formed [16,17].

The results show that the chemical reaction is the controlling stage for leaching copper from chalcopyrite concentrate. Figure 13 shows the linear regressions of the 9 tests. The equation of the chemical reaction as the controlling step is also shown as follows (Equation (7)).

$$kt = 1 - (1 - \chi)^{(1/3)} \tag{7}$$

where t is time and *k* is the apparent velocity constant.

Table 10 shows the apparent velocity constants (*k*) and the determination coefficient (R2) of the linear regression of the chemical reaction model for all 10 tests.


**Table 10.** Results for the linear regression of the chemical reaction model of all 10 tests.

To calculate the activation energy, logarithms were applied to the Arrhenius equation (Equation (8)) to reformulate it as a linear equation. Accordingly, the logarithm of the apparent velocity constants versus the inverse of temperature of tests P3, P8, P9 and P10 is shown in Figure 14.

$$\text{Logk} = \text{LogA} - \log\left(\text{E/R}\right)^{\left(1/\text{T}\right)}\tag{8}$$

An activation energy of 61.93 kJ/mol was determined from the slope of the straight line in Figure 14. According to Habashi (1999), a chemically-controlled process is usually greater than 41.8 kJ/mol [18].

**Figure 13.** Linear regression of the chemical reaction model.

**Figure 14.** Logarithm of the apparent velocity constant versus the inverse of temperature.

In order to compare the activation energy with similar processes, Table 11 shows the results reported by some authors in literature. It can be observed that the activation energy for processes that use sulfuric acid, ferric ions and/or oxygen is similar to the obtained in this work that use sulfuric acid and oxygen. In the work reported by Padilla et al. (2008), they use also only sulfuric acid and oxygen in a pressure reactor and a copper extraction of 95% was obtained at an oxygen pressure of 5 kg/cm2, 125 ◦C in 4 h [19]. In our work, we obtained 98% Cu extraction under less extreme conditions: oxygen pressure of 1 kg/cm2, 100 C and only 3 h.



Padilla et al. (2008) required a higher activation energy than the present work; this is due to the design of reactor used. The reactor used in the present work promotes a high interaction between the solid-gas-liquid phases, improving the mass transport at the gas-solid interface. Thus, the activation energy required for the leaching of chalcopyrite decreases.

## **4. Conclusions**

In the present study, the copper leaching of chalcopyrite concentrate in a 30-L batch reactor was described. The experimental results showed that it is possible to extract 98% of copper in only 3 h. This result indicates a fast process compared with others reported in literature.

The best result (98% in 3 h) was obtained under the following reaction conditions: 130 g/L of initial sulfuric acid concentration, temperature of 100 ◦C, oxygen pressure of 1 kg/cm2, solid concentration of 100 g/L and concentrate particle size of −104 + 75 μm.

The copper leaching is controlled chemically. Then, the elemental sulfur layer exposed on the unreacted particles of chalcopyrite does not interfere with the mass transport or the interactions between phases.

A statistical analysis showed that temperature is the most important variable influencing the extraction of copper and oxygen consumption. A temperature of at least 92 ◦C (61.93 kJ/mol) is necessary to activate the decomposition of chalcopyrite.

The initial sulfuric acid concentration must also be considered as an important variable from an economic perspective. An excess of sulfuric acid will increase the neutralizing agent in the posterior stages of the leaching process, whereas a lack of sulfuric acid could result in the precipitation of iron as plumbojarosite, and could therefore create difficulties in the recovery of valuable metals at later stages.

**Author Contributions:** Conceptualization, J.C. and R.C.; Formal analysis, J.C.; Investigation, J.C., J.P., J.V. and R.C.; Methodology, J.C., J.V. and R.C.; Project administration, I.A.; Resources, J.P. and I.A.; Supervision, J.P., R.C. and I.A.; Validation, J.C. and I.A.; Writing—original draft, J.C.; Writing—review & editing, J.V. and R.C.

**Acknowledgments:** The authors gratefully acknowledge the technical support of the National Institute of Technology of México and Servicios Especializados Peñoles S.A. de C.V. for the financial support.

**Conflicts of Interest:** The authors declare no conflict of interest.

## **References**


© 2019 by the authors. Licensee MDPI, Basel, Switzerland. This article is an open access article distributed under the terms and conditions of the Creative Commons Attribution (CC BY) license (http://creativecommons.org/licenses/by/4.0/).

## *Article* **A Kinetic Study on the Preparation of AlNi Alloys by Aluminothermic Reduction of NiO Powders**

## **Cesar Silva Beltran 1, Alfredo Flores Valdes 1, Jesús Torres Torres <sup>1</sup> and Rocio Ochoa Palacios 2,\***


Received: 9 July 2018; Accepted: 31 July 2018; Published: 28 August 2018

**Abstract:** In this work, the experimental results obtained during the preparation of Al-Ni and Al-Ni-Mg alloys using the aluminothermic reduction of NiO by submerged powder injection, assisted with mechanical agitation are presented and discussed. The analyzed variables were melt temperature, agitation speed, and initial magnesium concentration in the molten alloy. For some of the experiments performed, it was found that the Ni concentration increased from 0 to about 3 wt-% after 90 min of treatment at constant temperature and constant agitation speed. In order to determine the values of the kinetic parameters of interest, such as the activation energy and the rate constants, the values of the results obtained were fitted to the kinetic formulae available. Moreover, the kinetics of the reaction were found to be governed by the diffusion of Al and Mg to the NiO boundary layer, where MgAl2O4 or Al2O3 were formed as the main reaction products. Finally, from a thermodynamic study of the system, the main reactions that took place are explained.

**Keywords:** Al-Ni alloys; aluminothermic reactions; reaction rate; Al master alloys; kinetics

## **1. Introduction**

Aluminum-nickel alloys are widely used in the manufacturing of parts for the automotive industry, particularly those with a high nickel content, around 3 wt-%, and in cutting tools with TiAl because of their high resistance and low corrosion [1–4]. It is expected that in the upcoming years, the production of this type of alloy will increase worldwide, to reach a level of production higher than 400,000 tons per year. The possibility of preparing Al-Ni and Al-Ni-Mg alloys from the aluminothermic reduction of NiO has been proposed, using the submerged injection of powders as a technique for incorporating the nickel oxide particles.

In this sense, the first applications of the aluminothermic reduction of oxides were in the preparation of molten iron streams for filling the enclosures of exposed surfaces, where the slag rich in Al2O3 could be separated easily because of its lower density relative to that of iron [5]. Currently, an aluminothermic reduction is used for the preparation of aluminum master alloys, such as Al-Sr-Mg, Al-Ce-Mg, Al-Zn, Al-Mn, Al-Zr, Al-Mg-Fe-Cr, Al-Li, and so on. [6–15]. Nowadays, aluminothermic reduction is the most common practice used in the production of both kinds of products, pure metals or alloys [16–19]. However, only recently has the understanding of the kinetics of reaction been addressed for the systems of reaction, as every system is quite different.

The reduction reaction of nickel oxide by molten aluminum has a considerably negative value for the standard Gibbs free energy of reaction at operating temperatures, as is shown by the next reaction:

$$\text{3MO} + \text{2Al} = \text{Al}\_2\text{O}\_3 + \text{3M} \tag{1}$$

In the specific case of NiO, it is possible to accelerate the reduction reaction of nickel oxide with magnesium, thereby increasing its concentration in the molten aluminum during the aluminothermic process. In this case, the aluminothermic reaction occurring at 1073 K could be as follows:

$$2\text{Al} + \text{Mg} + 4\text{NiO} = \text{MgAl}\_2\text{O}\_4 + 4\text{Ni}, \ \Delta G^{\circ}\_{\text{1073 K}} = -869.48\text{ kJ} \tag{2}$$

Magnesium enhances the reaction rate, not only through its effect on the chemical reactivity of the molten bath, but also through its effect on the surface tension of molten aluminum, improving the wettability between the NiO particles and the molten aluminum solution [20].

On the other hand, it is broadly known that in the last 30 years, a lot of polluting waste from secondary spent batteries has been generated, of which NiO and nickel hydroxide (NiOH) are the main components. As such, it is attractive to investigate the aluminothermic reduction rate of NiO to formulate and propose alternatives for the use of materials that come from the recycling industry, and that can have a high added value, such as the preparation of Al-Ni or Al-Ni-Mg type alloys. In this sense, the main purpose of this paper is to present and discuss the results of an investigation aimed at understanding the role of temperature, initial magnesium concentration in molten aluminum, and the degree of agitation during the preparation of Al-Ni and Al-Ni-Mg alloys using a submerged powder injection of NiO powders, assisted by mechanical agitation. The results mainly focus on measuring the reaction rate and values of some of the kinetic parameters of interest, such as the rate constants and activation energy values of the process, where describing the mechanism of the reaction is of paramount importance.

## **2. Materials and Methods**

A 70 KW Power Track 75–30 electromagnetic medium frequency induction furnace, equipped with a silicon carbide crucible with a capacity of 10 kg of molten aluminum, was used as the melting unit. The powder injection equipment, whose scheme is presented in Figure 1, allowed for the continuous and controlled feeding of the solid material (NiO) through an inert carrier gas, which in this investigation, was high purity argon (99.99%). The design of the powder injection equipment used for the tests basically consisted of a pressurized chamber, inside which was a cylindrical container to deposit the reactive powder. This container was provided with a worm on the inside, which was rotated by an electric motor connected to a voltage controller located on the control board. This allowed us to vary the speed of rotation of the worm, and therefore, control the speed of the feeding of the powder.

The mixture of powder/gas falls directly into a funnel that connected the feeding line to the outlet lance of the pressurized system, being dragged thereto by means of the carrier gas. In this way, a powder/gas mixture is conducted towards the melting unit, where it is injected inside the molten metal through a graphite lance. This lance is located at a depth of 85% of the height of the molten metal. The location of the injection lance with respect to the geometry of the furnace, plays a very important role in the reaction kinetics, facilitating the discharge of the powder, therefore assuring its contact with the molten metal solution. The dimensions of this lance were as follows: length, 50 cm; external diameter, 5.08 cm; and diameter of the internal hole, 3.54 cm. To promote a better agitation inside the furnace, a graphite agitator was built, which was assembled to a mechanical drill and placed in the center of the bath; in this way, the agitation was constant and vigorous, as, in addition to the agitator and the gas injected, the Eddy electromagnetic waves generated by the induction coil helped to reach a greater efficiency in the mixing conditions. An insulating lid was attached to the furnace to minimize the inlet of air to the furnace. This stage had an entry for the mechanical agitator and another for the injection lance.

**Figure 1.** Schematic illustration of the experimental set-up employed in the experiments of the submerged powder injection with mechanical agitation.

Once the lid and the agitator were placed, the temperature and the revolutions per minute were adjusted while the lance was inserted in the molten aluminum to later adjust the powder feeding speed, thus initiating the treatment. We conducted 27 experiments, where the effects of temperature, percentage of magnesium in the alloy, and agitation speed in the molten bath were investigated. The operational parameters and their values were selected from preliminary works [8–12], where the kinetics of the aluminothermic reactions have been studied, not only for attaining reliable reaction efficiencies, but also to avoid metal losses during treatment. Initially, a factorial design was proposed to reduce the experimental errors. The factorial design was comprised of three factors and three levels, as is shown in Table 1. The number of experiments was performed according to the data in Table 1 and Equation (3).


**Table 1.** Parameters and levels of the initial experimental design.

$$N = a \times b \times c \times n \tag{3}$$

where *N =* numbers of experiments; *a =* 3 (factor A levels) A level; *b* = 3 (factor B levels) B level; *c* = 3 (factor C levels) C level; and *n* = 1 (number of replicas).

According to this, the total number of experiments (*Ei*) is shown in Table 2. To calculate the number of reagents to be injected, the amount of aluminum, the nickel amount released from NiO, and the stoichiometry given by Reaction (2) were considered. Therefore, the amount of theoretical NiO that could feasibly be added to the molten bath was calculated as 250 g, also by considering the magnesium concentration in the corresponding trial tests that required it. The target Ni concentration was 3 wt-%. The parameters that remained constant were the initial quantity of molten aluminum, 5 kg; powder injection rate, 250 g min−<sup>1</sup> of NiO; argon injection rate, 5 L min−1; and particle size of the NiO powders, to the order of a 4 μm average.


**Table 2.** Design of total experiments (*Ei,j,k*). *i*—temperature (K); *j*—initial Mg concentration (wt-%); *k*—agitation speed (rpm).

Samples from the molten bath were taken every 10 min, attaining up to 90 min of treatment. These samples were poured into a metallic mold and marked accordingly for chemical analysis determinations and microscopic observations, using both optical and scanning electron microscopy (SEM). The response variables were the actual nickel and magnesium concentrations, and the amount of Ni or Ni-Mg rich phases as a function of the injection time. A chemical analysis was carried out using the spark emission spectrometry technique. The chemical composition of the aluminum alloy is given in Table 3, while the chemical composition of the NiO powders is given Table 4.

**Table 3.** Chemical composition of the aluminum alloy (wt-%).




At the end of each experiment, the samples from the produced slags were taken for characterization by X-ray diffraction. To determine the Standard Gibbs free energy of reaction, the software HSC [21] was used, while for the drawing diagrams of the phase stability as a function of the temperature of the reaction, the software Factsage Chemistry Software was used [22].

## **3. Results and Discussion**

## *3.1. Experimental Results*

Figures 2 and 3 show that the Ni concentration increased in the molten alloys as a function of the injection time for the temperatures and initial magnesium concentrations indicated, at the agitation speed of 300 rpm. From these graphs, the nickel concentration reached the highest value at higher magnesium concentrations, at the temperature of 1123 K.

**Figure 2.** Increase in nickel concentration as a function of injection time at the indicated temperatures, for 0 wt-% Mg, and a constant agitation speed of 300 rpm.

**Figure 3.** Increase in the nickel concentration as a function of injection time at the indicated temperatures for 3 wt-% Mg content in the alloy and at a constant agitation speed of 300 rpm.

Figure 4 shows the increase in the Ni concentration as a function of the injection time for the agitation speeds considered, at the constant initial Mg concentration and temperature indicated. From this figure, it can be observed how the agitation speed plays an important role in the aluminothermic reduction of the NiO particles, which was due to the improved mass transfer conditions attained at higher values of this parameter. Indirectly, it also indicated the diffusive nature of the process taking place.

**Figure 4.** Increase in the nickel concentration as a function of injection time, at the indicated agitation speeds, for an initial Mg concentration of 3 wt-%, and a temperature of 1123 K.

From the explanations of Guedes et al. [23], the increase in the nickel concentration as a function of the injection time, as shown in Figures 2 and 3, can be explained as follows. When magnesium is not present in the molten alloy, a layer of Al2O3 is formed at the beginning of the NiO reduction reaction. The growth of this layer minimizes the contact between the molten aluminum and the reaction front, acting as a barrier between both the aluminum and nickel oxide, which prevents the reaction given by Equation (2) to be completed. However, the Al2O3 layer is not completely impervious, thus having some porosity through which the aluminum can diffuse to the boundary layer, where it reaches the surface of the NiO particles. Magnesium improves the reaction rate in two ways. At 973 K, it reduces the surface tension of the molten aluminum from 0.91 (with no Mg present) to 0.66 N m−<sup>1</sup> for a magnesium concentration equal to 1 wt-% [24]. This facilitates the wettability of the NiO particles in the molten phase. On the other hand, Mg increases the reactivity of the bath, because this element also reduces the NiO to 973 K, according to Reaction (2), where the greater negative value of its Gibbs free energy indicates the spontaneity of this reaction.

During their experiments on the synthesis of the Al/Al2O3 composites, Pai and Ray [25] found that MgO could be formed from the reduction of Al2O3 by the magnesium contained in the alloy. This reduction reaction can also occur in the case of the NiO reduction reaction given by Equation (2). In turn, MgO and Al2O3 can react during solidification to form MgAl2O4 as the final reaction product, according to the following reaction between pure compounds, whose Δ*G*◦ is calculated at 1073 K.

$$\text{MgO} + \text{Al}\_2\text{O}\_3 \rightarrow \text{MgAl}\_2\text{O}\_4 \; \text{Al}^\circ \; \_{1073\text{ K}} = -92.02\text{ kJ} \tag{4}$$

The successive formation of MgO and Al2O3 causes expansion and contraction during the reaction, which results in breaking the crystals into numerous smaller crystals, thus producing the porosity required for the diffusion of aluminum, magnesium onwards, or nickel backwards.

Figure 5 shows the variation in the magnesium concentration as a function of the NiO injection time, and the concentration of magnesium decreased continuously during the injection, reaching the lower value of 1.98 wt-% at the end of the experiments for the given conditions established.

**Figure 5.** Decrease of the magnesium concentration in molten aluminum as a function of injection time at the indicated temperatures at the constant agitation speed of 300 rpm.

The micrographs in Figure 6 show the evolution of the microstructure of the aluminothermic reduction reaction of the NiO powders as a function of the indicated injection times from samples obtained at 1123 K, at a constant agitation speed of 300 rpm. In these micrographs, it was observed that the amount of the nickel-rich intermetallic phase increased as the treatment time increased, therefore the amount of nickel incorporated in the molten alloy increased with increasing time.

Figure 7 shows the increase in the amount of Ni-rich intermetallic phase as a function of the injection time, hence, the amount of nickel incorporated was increased in the molten alloy by the aluminothermic reduction reaction of the NiO powders.

**Figure 6.** Photomicrographs of Al-Ni samples obtained at 1123 K, as a function of injection time, where (**a**) 0 min, (**b**) 20 min, (**c**) 40 min, and (**d**) 90 min.

**Figure 7.** Increase of the Al3Ni intermetallic phase as a function of injection time for the temperatures indicated at the constant agitation speed of 300 rpm.

The micrographs in Figure 8 show the evolution of the thickness of the layers of the NiO particles reacted as a function of the reaction time, taken from the molten metal at a temperature of 1123 K for an agitation speed of 300 rpm. After 30 min, the particle remained unreacted (Figure 8a), and after 40 min of addition, it began to form an Al2O3 layer around the NiO particle (Figure 8b). This layer grew progressively until the reaction stopped because of a change in the mechanism that controls the reaction (Figure 8c,d). After 70 min of contact (Figure 8e), the Al2O3 layer began to disappear. In the micrographs, it can be seen that after 90 min of reaction, the Al2O3 layer completely disappeared, leaving only the solid particle and the Al3Ni intermetallic particles around them.

**Figure 8.** SEM images of partially reacted NiO particles obtained from molten aluminum at 1123 K at a stirring speed of 300 rpm for the indicated times of (**a**) 30 min, (**b**) 40 min, (**c**) 50 min, (**d**) 60 min, (**e**) 70 min, and (**f**) 90 min.

The particle shown in Figure 8d was analyzed separately by energy dispersion spectroscopy in the SEM, and the corresponding energy dispersive X-ray spectroscopy (EDS) patterns, together with the micrographs showing the microareas from where they were obtained, are shown in Figure 9. The microanalysis measurements show that the nucleus of NiO was surrounded by an Mg and Al-rich phase, marked as (2) and (3). The semi-quantitative chemical composition of the layer of the reaction product, as shown in Table 5, was observed as the phase formed was spinel (MgAl2O4). These EDS results suggest that the NiO was reduced by aluminum and magnesium, forming intermediate reaction products such as the spinel phase, while the Ni-rich phase corresponded to Al3Ni [26].

**Figure 9.** SEM images of the partially reacted NiO particle shown in Figure 8d, and the corresponding EDS patterns of the microareas of the (**1**) center, (**2**) outside layer, and (**3**) MgAl2O4.

**Table 5.** EDS chemical analysis of the partially reacted NiO particle shown in Figure 8d.


Figure 10 shows the X-ray powder diffraction (XRD) pattern corresponding to the slag obtained after the end of experiment E27, where there was the presence of MgO, MgAl2O4, aluminum, and NiO. The aluminum came from the molten alloy, as the sample was taken from a pasty area formed between the slag and molten aluminum.

The MgO and MgAl2O4 in the slag suggests that both aluminum and magnesium reduced the nickel oxide by a metallothermic mechanism. The NiO came from the partially reacted particles. The analysis of the slag helped to corroborate that the main aluminothermic reduction reaction was that given by Equation (2), where the Ni was obtained from NiO dissolved in the molten alloy to solidify as Al3Ni particles, which are shown in Figure 5.

**Figure 10.** XRD pattern of the slag samples for an alloy where the initial magnesium was 3 wt-% at a stirring speed of 300 rpm and a temperature of 1123 K.

## *3.2. Thermodynamic Consideration for the Al-NiO-Mg System*

The aluminum alloys obtained in this work contained between 2 to 3 wt-% of both Mg and Ni, therefore, a thermodynamic explanation would help to understand the expected results. During the aluminothermic reduction process, the Al, Mg, and NiO species are involved, and react to obtain different reaction products. The software FactSage allowed us to consider ideal solutions for the Al-Mg-Ni system in the solid and liquid states. The reactants that were introduced into the software were Al, Mg, and Ni as pure species. The balance of the system was computerized for the temperature range from 373 to 1273 K, where the activities considered for Al, Ni, and Mg were equal to 1. Figure 11 shows the results of the software in the Al-Mg-Ni equilibrium system. The phases that were present below the melting temperature of aluminum were Al(s), MgO, and Al3Ni in their stable forms, while at temperatures higher than 933 K, the Al and Mg were in a liquid state. These results define the reactions involved during the reduction of NiO powders by Al-Mg(l). The formation of MgO is thermodynamically feasible from room temperature up to the melting temperature of aluminum. However, the diagram drawn by the software was preliminary as the kinetic conditions change continuously during the reaction.

**Figure 11.** Equilibrium diagram of species for the Al-Mg-NiO system in the temperature range from 373 to 1873 K.

The reactions are presented and the values of Δ*G*◦ at 1123 K were obtained from the software, expressed per mole of oxygen.

$$4/3\text{Al} + \text{O}\_2 = 2/3\text{Al}\_2\text{O}\_3 \text{ } \text{ $\stackrel{\circ}{\text{C}}\_{1123\text{ K}} = -882.46\text{ kJ}$ }\tag{5}$$

Reaction (5) can occur between aluminum and oxygen from the environment with pressures of 1 atm and elevated temperatures, forming Al2O3. The next reaction can occur between Al2O3 with magnesium and oxygen to form the spinel MgAl2O4, which is possible with amounts higher than 0.05% Mg.

$$\text{Al}\_2\text{O}\_3 + \text{Mg} + 1/2\text{O}\_2 = \text{MgAl}\_2\text{O}\_4 \text{ } \Delta \stackrel{\circ}{G}\_{1123\text{ K}} = -1036.56\text{ kJ} \tag{6}$$

Reactions (5) and (6) must be produced at the aluminum/slag interface, forming a protective layer of aluminum and magnesium oxides in the molten bath.

The reaction between Al and NiO to form Al2O3 while releasing Ni is quite possible because of the action masses law, as aluminum is the element present in greater quantity in the system. The reaction is that given by Equation (1). However, the graphs in Figure 2 show that in the experiments performed with pure aluminum, the concentration of nickel attained was rather low. It was shown that Al2O3 [27] in a wide range of temperatures was not wettable by aluminum, because it has a contact angle value >90.5◦ [28].

Therefore, it is necessary to decrease the contact angle of Al2O3 by molten aluminum. According to the literature [29], one way to change the contact angle is to lower the surface energy of molten aluminum by adding Mg, as this is one of the elements that has this effect.

In the experiments that were carried out, it was observed that the magnesium content decreased with the increased time, as is shown in Figure 5. This was attributed not only to the effect of this element on the contact angle, but also to the high reactivity that magnesium possesses, according to the following reaction:

$$\text{NiO} + \text{Mg} = \text{MgO} + \text{Ni}, \ \text{Al}^{\circ}\_{\text{ }1123 \text{ K}} = -342.83 \text{ kJ} \tag{7}$$

The global chemical reaction taking place is that given by Equation (2). The following reaction is possible because of the greater chemical reactivity of Mg than Al:

$$\text{Al}\_2\text{O}\_3 + 3\text{Mg} = 3\text{MgO} + 2\text{Al}, \text{ Al}^\circ\_{1123\text{ K}} = -117.69\text{ kJ} \tag{8}$$

The formation of MgO, a phase present in the slag produced in this work for aluminum alloys with initial Mg concentrations above 2 wt-%, proceeded from the following reaction:

$$2\text{Mg} + \text{O}\_2 = 2\text{MgO}, \ \text{\AA}\text{\AA}^\circ \text{\AA}\_{1123\text{ K}} = -970.09\text{ kJ} \tag{9}$$

In turn, Al2O3 and MgO react according to Equation (4) to form the so-called spinel, MgAl2O4. All of the reaction products described by the reactions given by Equations (1), (2), (4), (5), or (7)–(9) were found in the slags obtained from the NiO reduction process studied in this work.

## *3.3. Kinetic Measurements*

The experimental evidence strongly indicated that the reaction rate in the studied process was controlled by the diffusion of chemical species, mainly Al or Mg, to the boundary layer, where they reacted with NiO. Consequently, the experimental information obtained at each temperature, for alloys with and without magnesium, was analyzed using the kinetic formulas presented by Brown [30]. However, to adapt to the experimental information of the available kinetic models, it was assumed that the observed nickel concentration corresponded to that of the complete reaction, and therefore, the final reacted fraction was equal to 1. First of all, it is necessary to state that for the present case, the reaction rate is defined as the ratio of the reacted fraction (*α*) against time (*t*), given by the following:

$$
\text{reaction rate} = \frac{d\alpha}{dt}\tag{10}
$$

As an associated magnitude to the progress of the aluminothermic reaction, *α* is numerically equal to the reacted fraction attained in a given period, called actual time. Therefore, reacted fraction, in our case *αNi*, is given by next equation:

$$\alpha\_{Ni} = \frac{w\_o - w\_i}{w\_o - w\_f} \tag{11}$$

where *w*o, *wi*, and *wf* are the *Ni* concentrations measured in the alloy at the beginning of the injection time, the actual concentration, and the final concentration, respectively.

Figures 12–14 show the change in the reacted fraction as a function of time, for the different set of parameters studied, resulted from the application of Equation (11) to the experimental results obtained. In general, it can be observed from these graphs that the aluminothermic reduction reaction was not completed for any value of the parameters studied (i.e., temperature, initial magnesium concentration, or agitation speed), owing to the diffusive nature of the reaction, as it will be explained later.

**Figure 12.** Reacted fraction of Ni as a function of time for the temperatures indicated, for a constant initial concentration of Mg of 3 wt-%, at the constant agitation speed of 300 rpm.

**Figure 13.** Reacted fraction of Ni as a function of time for the initial magnesium concentrations indicated, at the constant temperature of 1123 K, and constant agitation speed of 300 rpm.

**Figure 14.** Reacted fraction of Ni as a function of time for the agitation speed indicated, for a constant initial concentration of Mg of 3 wt-%, at the constant temperature of 1123 K.

In this sense, it was found that the kinetic information was well suited for the diffusion model equation (D1), as shown in Figure 15, which corresponded to the experiments where alloys with magnesium were used at an agitation speed of 300 rpm. In turn, Figure 16 shows the experimental results at different temperatures. The equation of the model is as follows:

$$(1 - \mathfrak{a})\ln(1 - \mathfrak{a}) + \mathfrak{a} \tag{12}$$

where *α* is the reacted fraction.

**Figure 15.** Experimental results and prediction of model D1 at the temperature of 1123 K for an alloy with 3 wt-% Mg at 300 rpm.

**Figure 16.** Experimental results and prediction of the D1 model at different temperatures for alloys with an initial Mg concentration of 3 wt-%, at 300 rpm.

The linear regression of *ln* (*k*) against 1/*T* was plotted to obtain the values of the activation energy and the rate constants using the Arrhenius equation given below:

$$k = k\_0 e^{-E/RT} \tag{13}$$

where *E* is the activation energy (J mol−1), and *R* is the universal gas constant (8.314 J mol−<sup>1</sup> K). Figure 17 shows the linear dependence of a graph of *ln* (*k*) against 1/*T* for the aluminothermic reduction of NiO for alloys with magnesium (a) and without magnesium (b). From the slope values of these graphs, the values of the activation energies were determined as *E* = 35.75 KJ mol−<sup>1</sup> for the alloys

without magnesium and *E* = 15.80 KJ mol−<sup>1</sup> for the alloys with an initial magnesium concentration of 3 wt-% Mg.

**Figure 17.** Graphs of *ln* (*k*) against 1/*T* for the determination of the activation energy for the recovery of nickel with magnesium were 0 and 3 wt-%.

From the values of the kinetic parameters determined in this section, it can be stated that, as expected, the higher values of the rate constants indicated that the reaction rate was improved by the presence of magnesium in the alloys, at least initially. This occurred because, for reaction times greater than 70 min, the reaction tends to stop. According to the observations in the micrographs of the partially reacted particles, it can be stated that this occurred through the formation of a thin layer of reaction products around the NiO particles. In addition, the activation energy for the aluminothermic reduction reaction of the NiO particles when using alloys with 3 wt-% by weight of magnesium, was lower than that of the experiments when no Mg was added.

## *3.4. Mechanism of Reaction*

Regarding the reaction mechanisms that operate during the aluminothermic reduction reaction of NiO powders, the most precise explanation is that given by Zhong et al. [31], who stated that the formation of MgAl2O4 depends strongly on the initial concentration of magnesium, as has been previously established thermodynamically and has been experimentally proven in this work. Therefore, for initial magnesium concentrations of around 1 wt-%, the stable phase is MgAl2O4. Mcleod and Gabryel [32] established that the presence of MgO occurred because of the high initial concentration of magnesium (>3 wt-%). However, this concentration was easily reached at the NiO interface, especially in the initial stages after the addition of the particles, which ensured the formation of a significant amount of MgO. On the other hand, most of the MgO was consumed by the MgAl2O4 formation reaction given by Equation (4). Molins et al. [33] studied the interfacial reaction between a molten AlMg alloy and Al2O3 fibers, where it was shown that the MgO nuclei remained small and formed thin layers of ~10 μm thick. Therefore, the diffusion of magnesium occurred through the grain boundaries of the MgO particles by means of an infiltration mechanism. The growth process of the MgO layers continued until the grains around the matrix/reaction zone interface were large enough to close the intergranular spaces. The Gibbs free energy reaction of Equation (2) indicated that the reduction of NiO by aluminum occurred because Al can diffuse through the spaces left by the backwards diffusion of Ni. The diffusion of nickel occurred through the grain boundaries of MgO and Al2O3, dissolving

in the molten aluminum when it reached the outer interface. Upon solidification, the nickel was rejected from the solid solution, forming Al3Ni crystals around the partially reacted NiO particles. When the stoichiometric amounts of MgO and Al2O3 formed and equilibrium conditions occurred, the reaction given by Equation (2) was carried out. In this way, the formation of MgO and Al2O3 and the interdiffusion of Al, Mg, and Ni occurred simultaneously. After the magnesium concentration in the molten alloy fell below 1.98 wt-%, the aluminum continued to react with the NiO cores to form additional Al2O3 at a constant growth rate controlled by the diffusion of the aluminum through the layers of the reaction products. On the other hand, the small amount of MgO that could be formed by the magnesium reaction with NiO was dissolved in the Al2O3 phase. The reactions between the molten aluminum and magnesium dissolved with solid NiO require the diffusion of the atoms, although the diffusion of oxygen is negligible due to its large ionic size.

Therefore, the interdiffusion of magnesium and aluminum is kinetically possible through the network in the NiO structure of the O atoms, or through the nickel that results from the decomposition of the NiO. When the magnesium atoms occupy the interstitial sites within the NiO network, the position of the oxygen atoms must be adjusted to maintain electrical neutrality and to decrease the distortion energy of the network. Such reactions were given by the following:

$$\text{Mg}^{0} - 2\text{e}^{-} = \text{Mg}^{+2} \tag{14}$$

$$2\left(\text{Al}^0 - 3\text{e}^0\right) = 2\text{Al}^{+3} \tag{15}$$

$$4\left(\text{Ni}^{+3} - 3\text{e}^{-}\right) = 4\text{Ni}^{0}\tag{16}$$

The result is a change in the crystalline structure. The process involved in this phenomenon is governed by the values of the chemical potential of the partial reactions given, in this case, by Reactions (1) and (2). As MgAl2O4 is the final reaction product at room temperature, the reaction given by the global Equation (2) can be accepted, as this reaction satisfies the described mechanism. A necessary kinetic condition for the diffusion of chemical species is that the porosity remains in the layers of the reaction products. For similar reduction chemical reactions in molten aluminum, Zhong et al. [31] determined that the formation of MgO on SiO2 particles involved a volume contraction of 13.6%, while the formation of MgAl2O4 on SiO2 particles caused a volume contraction of 27%. Due to these changes in volume, the newly formed phases instantly broke their junctions with the original particles and transformed into thousands of small crystals. This transformation produces the necessary voids for the diffusion of the chemical species. However, for the last stages of the reduction process, the reduction mechanism changed abruptly, as the concentration of magnesium in the reaction interface decreased to below 1.98 wt-%. At this time, the MgO that was formed dissolved in Al2O3. Then, these Al2O3 crystals grew continuously until they reached a micrometric size, which resulted in the blocking of spaces for diffusion and the chemical reaction begins to stop. In this last stage, the efficiency of the reaction decreased sharply. The above explanations are important from the technological point of view, as unreacted NiO particles can be trapped in the molten metal as inclusions, and thus also affect the efficiency of the reaction. Of course, the kinetics of the reaction can be accelerated by further decreasing the size of the NiO particles or imposing the mixing conditions in the turbulent flow regime by using Reynolds numbers above 4.5 × <sup>10</sup><sup>3</sup> to break the layers of the reaction products once they have formed [34].

Figure 18 shows a schematic representation of the proposed mechanism of reaction, which described the different stages and the reactions among the participant chemical species, corresponding to a situation where the temperature was constant at 1123 K, initial magnesium concentration was constant at 3 wt-%, and agitation speed is constant at 300 rpm.

**Figure 18.** Scheme of the mechanism of reaction proposed for the permanent contact reaction between dissolved magnesium in molten aluminum and NiiO particles; (**a**) instantaneous formation of MgO and Al2O3, 10 min; (**b**) MgAl2O4 formation, 40 min; (**c**) nickel diffusion to the molten alloy, 70 min; (**d**) and consolidation of the layers of reaction products consisting of MgAl2O4 separated by layers of the Al2O3 phase, 90 min.

## **4. Conclusions**


5. With respect to the reaction mechanism, it was found that the step that controlled the overall chemical reaction was the diffusion of the Al and Mg atoms to the boundary layer, where they reacted with NiO particles, releasing Ni and forming Al2O3 and MgO as the reaction products. In turn, these compounds formed MgAl2O4 during cooling. The formation and breaking of MgAl2O4 into many crystals ensured the porosity required for the diffusion of the chemical species involved.

## **Author Contributions:** Investigation, C.S.B. and R.O.P.; supervision, A.F.V. and J.T.T.

**Funding:** This research was funded by Fundicion J.V. with the project 217843 of the program Stimulus to the Investigation of CONACYT Mexico.

**Acknowledgments:** The authors wish to acknowledge the financial support of Fundicion J.V., project 217843 of the program Stimulus to the Investigation of CONACYT México, as well as for the cast shop facilities provided through this project.

**Conflicts of Interest:** The authors declare no conflicts of interest.

## **References**


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