**4. Process description**

In this work, we propose a process for leaching of the main constitutes of the SPL waste by H2SO4 solution. The combination of **Figures 1, 2** and **3** constitute the process flow diagram (PFD) of the proposed leaching process. Note: The numbers in red color beside the stream numbers on these figures, are the stream input temperature (30°C) or the calculated temperature using heat of mixing and reaction thermochemical data along with the energy balance equations. Most of the acid leaching reactions are exothermic (-ΔHR) except those appearing in bold numbers in the ΔHR column of **Table A.1** in particular.

#### **Figure 1.**

The collected SPL waste first passes through crushing and grinding steps. The resulting SPL fines are fed to an agitated semi-batch reactor filled with a preprepared H2SO4 solution. To ensure that all the SPL particles are sufficiently exposed to the solution, a 2.5 M H2SO4 (with 5 wt% excess) is used along with a recommended L/S ratio of 2.52 liters of H2SO4 acid solution per kg of SPL [19]. The reactor contents should be kept under agitation for 2–4 h. A 40,000 tons of SPL is assumed to be processed annually (or 5930 kg/h based on a stream factor of 0.77). However, a total of 220 working days per year (batch-wise operation, 22 working days per month, and allowing 2 months for shutdown and maintenance, i.e. stream factor = 0.6) is suggested elsewhere [19].

Considering the composition ranges of the SPL main constituents reported in [12] and presented in **Table 2**, the composition, the mass and molar flow rates based on the SPL upper composition limit are given in **Table 11**.

The products generated during processing are classified into three categories or streams: (1) gaseous stream (HCN, HF and CO2), (2) insoluble products stream (graphite, gypsum and SiO2), and (3) soluble products stream (aluminum fluorides and sodium salts, mainly, Na2SO4). Details on processing of each of these streams are given below and demonstrated in **Figures 1, 2** and **3** generated by the authors.

*A Zero-Waste Process for the Treatment of Spent Potliner (SPL) Waste DOI: http://dx.doi.org/10.5772/intechopen.99055*

#### **Figure 2.** *Process flow diagram and material balance for the SPL treatment … continued.*

1.During the leaching step, a gas stream (mainly, HCN, HF and CO2) leaves reactor R-101, cooled (not shown on the PFD) and then sent to a gas emissioncontrol scrubber (T-101) where the HCN gas is scrubbed by its reaction with a silver nitrate (AgNO3) solution sprayed at the top. See **Figure 1**. This reaction is spontaneous and exothermic. As a result, silver cyanide (AgCN) is produced according to Eq. (26). See reaction R8 in **Table A.1**.

$$\text{HCN(g)} + \text{Ag(NO}\_3\text{(aq)} \rightarrow \text{AgCN(s)} + \text{HNO}\_3\text{(aq)} \tag{26}$$

The AgCN is insoluble in water, but it is slightly soluble in aqueous HNO3. The AgCN, is separated from the aqueous solution via filter F-103. The AgCN salt is stable at ambient conditions and is very valuable in gold extraction. However, it is highly toxic by ingestion and its contact with skin and eyes can cause severe irritation. It has a LD50 oral (rat) of 123 mg/kg.

Note: It should be mentioned that no reaction will take place between aqueous AgNO3 used in Eq. (26) and HF(l), HF(g) or CO2, since these reactions are non-spontaneous at temperatures ≤90°C.

The HF can be recovered as a liquid from the HF-CO2 gas mixture by cooling/ condensation in E-101 to below its condensation temperature (at its partial pressure in the gas stream). The remaining gas from E-101 is sent to a CO2

#### **Figure 3.**

*Process flow diagram and material balance for the SPL treatment … concluded.*

recovery unit. The recovered HF liquid is pumped (P-101) where part of it is used within the process to ensure that all the remaining aluminum sulfate is converted to AlF3 (as explained below). The remaining part of the HF liquid can be sold as is or converted to potassium fluoride (KF); an important source of fluoride in many industries.

On the other hand, the normal boiling points of HF and HCN are 25.6°C and 19.5°C, respectively. Thus, one much better option (and much cheaper than scrubbing by AgNO3 solution) is the condensation of the HF gas followed by the condensation of HCN gas at their partial pressures in the gas phase stream leaving reactor R-101. This option avoids using the very expensive AgNO3 salt, but in this case, the condensed HCN must be destroyed by direct oxidation or it can be converted to a stable NaCN (soluble) salt by reacting HCN liquid with NaNO3 (very cheap). But still a reactor and a separator are


*A Zero-Waste Process for the Treatment of Spent Potliner (SPL) Waste DOI: http://dx.doi.org/10.5772/intechopen.99055*

#### **Table 11.**

*Normalized composition of the SPL main constituents used in this work.*

needed. In either case, the resulting gas stream needs to be sent to the CO2 recovery unit.

2.After completion of the leaching step, the slurry mixture is sent to filter F-101 where the insoluble solids (SiO2, graphite and gypsum) are separated from the aqueous solution containing soluble intermediate and final products (Na2SO4, AlF3 (and/or AlF2OH), remaining Al2(SO4)3, unreacted H2SO4, and water).

The insoluble solids stream is sent to reactor R-103 where the SiO2 is reacted with aqueous NaOH to produce soluble sodium silicate (Na2SiO3) according to reaction (27). See reaction R9 in **Table A.1**.

$$2\text{SiO}\_2(\text{s}) + 2\text{NaOH}(\text{aq}) \rightarrow \text{Na}\_2\text{SiO}\_3(\text{aq}) + \text{H}\_2\text{O} \tag{27}$$

which is then separated from the graphite-gypsum solid mixture via filter F-102. See **Figure 2**. The Na2SiO3 in the aqueous solution can then be saturated by evaporation and precipitated as Na2SiO3 crystals (not shown on the PFD).

The graphite and gypsum can be then separated from each other in a froth flotation unit (FF-101) where an oil (e.g. 1–10 wt% kerosene) in water is used, along with air bubbling and slow agitation. See **Figure 3**. The recommended particle size for froth flotation lies between +25 and 75 μm [42]. The hydrophobic graphite along with kerosene floats up as a froth while the hydrophilic gypsum along with water settles to the bottom of the unit. The graphite-kerosene stream is sent to filter F-106 to recover the graphite and recycle the kerosene back to the froth flotation unit. Similarly, the gypsum-water stream is sent to filter F-107 to recover the gypsum and recycle the water back to the froth flotation unit.

It should be mentioned that we have experimentally separated the graphite carbon from gypsum (using a kerosene/water volumetric ratio = 0.1 along with air bubbling at room temperature).

3.The aqueous phase from filter F-101 is cooled in E-102 and then sent to reactor R-102, where the remaining Al2(SO4)3 is converted to AlF3 (and/or AlF2OH) by its reaction with part of the recovered HF liquid, according to the relatively high spontaneous Eq. (28) (ΔGR = -196.65 kJ/mol). at 30°C. See reaction R10 in **Table A.1**.

$$\rm Al\_2(SO\_4)\_3(aq) + 6HF(l) \to 2AlF\_3(aq) + 3H\_2SO\_4 \tag{28}$$

Due to the presence of fluoride ions in R-102, the dominant crystal species will be AlF3. However, the reaction between Na2SO4 and HF(l) is much less competent than Eq. (28) since it is much less spontaneous (ΔGR = -32.7 kJ/mol). See reaction R1 in **Table A.1**.

In order to recover the AlF3 crystals, the contents of reactor R-102 are pumped through P-102 to the reactor-crystallizer RC-101, where the conditions required for AlF3 crystallization have to be established. A controlled amount of NaOH has to be added to neutralize most of the remaining H2SO4 according to Eq. (29). See reaction R11 in **Table A.1**.

$$\text{H}\_2\text{SO}\_4 + 2\text{NaOH} \rightarrow \text{Na}\_2\text{SO}\_4(\text{aq}) + 2\text{H}\_2\text{O} \tag{29}$$

and at the same time to maintain the solution in RC-101 at a pH of 4.5–5.5; required to saturate and precipitate AlF3 [19], noting that the solubility of AlF2OH (and AlF3) decreases with the increase of the pH.

Any AlF2OH produced can be easily converted to AlF3 by its reaction with some of the HF liquid recovered earlier, according to the spontaneous presented by Eq. (30). See reaction R12 in **Table A.1**.

$$\text{AlF}\_{\text{x}}(\text{OH})\_{3-\text{x}} + \text{x }\text{HF}(\text{l}) \rightarrow \text{AlF}\_{3} + \text{x }\text{H}\_{2}\text{O} \ (\text{x} = \text{1 or 2}) \tag{30}$$

or,

$$\text{AlF(OH)}\_{2} + 2\text{HF(l)} \rightarrow \text{AlF}\_{3} + 2\text{H}\_{2}\text{O} \text{ (for } \mathbf{x} = \mathbf{1}\text{)}\tag{31}$$

and

$$\text{AlF}\_2\text{OH} + \text{HF}(\text{l}) \rightarrow \text{AlF}\_3 + \text{H}\_2\text{O} \text{ (for } \text{x} = 2\text{)}\tag{32}$$

Thus, the reaction presented by Eq. (30) can be carried out before the addition of the NaOH solution.

The crystals produced in the reactor-crystallizer RC-101 are separated via filter F-104 as AlF3 cake. To remove the impurities from the AlF3, the stream needs to be washed with fresh water. The AlF3 is then dried, cooled and stored.

The filtrate leaving filter F-104 is sent to the evaporator-crystallizer EC-101, where the Na2SO4 solution is saturated by flash evaporation under vacuum and Na2SO4 is crystallized and separated via filter F-105. See **Figure 3**. The Na2SO4 crystals can be further dehydrated and dried before being stored.

Lastly, the water vapor leaving EC-101 is condensed in E-103 and collected for reuse within the process, along with other recovered water from the various streams of the above described process.

#### **5. Preliminary economic analysis**

A preliminary economic analysis has been made on the above proposed process (assuming a theoretical 100% conversion and/or recovery) following the guidelines of ref. [43]. The amounts and costs of raw materials used as well as the amounts and *A Zero-Waste Process for the Treatment of Spent Potliner (SPL) Waste DOI: http://dx.doi.org/10.5772/intechopen.99055*

market prices of the materials produced are listed in **Table 12**. The annual cost or price of a given material = amount (kg/h) x unit cost or price (\$/kg) x 6475.2 (h/year). The 6475.2 factor comes from 0.77 x 24 x 365. We made a preliminary design for the process equipment and estimated the fixed capital cost of the plant excluding land, FCIL, to be 27.32 M\$.

The number of operators per job was estimated based on Eq. (33):

$$\mathbf{N}\_{\rm OL} = \left(\mathbf{6.29} + \mathbf{31.7}\,\mathbf{P}^2 + \mathbf{0.23} \ast \mathbf{N}\_{\rm np}\right)^{0.5} \tag{33}$$

where P stand for particulate (solid) and Nnp for non-particulate (fluid) handling equipment (P = 1 for FF-101, Nnp = 15). The total number of operators required over the year = 4.47 NOL. The salary per operator was assumed to be \$49000.

The FCIL along with the estimated annual costs of labor COL, raw materials CRM, utilities CUT, and waste treatment CWT (given in **Table 13**) were used to calculate the cost of manufacturing excluding depreciation, COMd, according to Eq. (34):

$$\text{COMP}\_{\text{d}} = \text{0.18 FCI}\_{\text{L}} + \text{2.73 C}\_{\text{OL}} + \text{1.23} \left( \text{C}\_{\text{RM}} + \text{C}\_{\text{UT}} + \text{C}\_{\text{WT}} \right) \tag{34}$$

The calculated COMd = 21.73 M\$/year.

Now, assuming priceless produced HNO3, Na2SiO3, CO2 and output water, the income from main sales (revenue, R) was found to be 38.09 M\$/year. Also, since AgNO3 and AgCN are very expensive and sharply affect the profitably of the process, this option has been excluded in the economic analysis.


#### **Table 12.**

*Amounts of raw materials and products and their average prices [44].*


#### **Table 13.**

*Estimated individual operating costs and COMd.*

### *Waste Material Recycling in the Circular Economy - Challenges and Developments*


#### **Table 14.**

*Input data for discounted cumulative cash flow analysis.*

**Figure 4.** *Discounted cumulative cash flow diagram. (DCCF) for the above studied process.*

The input data used for generating the cumulative cash flow analysis are presented in **Table 14**. The discounted cumulative cash flow diagram for the above process analysis is shown in **Figure 4**. Following [43] economic analyses and using the data presented above, and assuming an interest rate of 10%, a tax rate of 20%, the calculated net present value, NPV = 42.24 M\$, the discounted payback period, DPBP = 2.38 years, and the discounted cash flow rate of return, DCFROR = 31.73%.

#### **6. Conclusions**

In this work an environmentally friendly process to recover the valuable elements contained in the SPL is presented and deeply analyzed. The decision to use H2SO4 as a leachant was justified through deep analysis. The proposed process along with the process flow diagram and complete material balance results have been explained and included.

*A Zero-Waste Process for the Treatment of Spent Potliner (SPL) Waste DOI: http://dx.doi.org/10.5772/intechopen.99055*

The recovered materials include graphite carbon, aluminum fluoride (AlF3), sodium sulfate (Na2SO4), and others when H2SO4 is used as the leaching agent. The level of emission of the hazardous gases such as HCN and HF are minimized. The recovered HF liquid is partially used within the process. The remaining HF can be used in production of potassium fluoride (KF). Also, CO2 gas can also be recovered from the process gas streams.

The economic analyses indicate that the process will be profitable under the conditions stated in this work. The process net present value, NPV = 42.24 M\$, the discounted payback period, DPBP = 2.38 years, and the discounted cash flow rate of return, DCFROR = 31.73%.
