METHOD FOR SULFUR REMOVAL FROM COAL FLY ASH

Abstract

A method of reducing sulfur concentration in fly ash, flue gas desulfurization (FGD) ash, and mixtures thereof by contacting the fly ash, FGD ash, or mixtures thereof with an aqueous acidic solution, for a time, at a temperature, and at a liquid-to-solid ratio wherein the sulfur concentration within the fly ash, FGD ash, or mixture thereof is reduced to no more than 5 wt % SO.sub.3 based on the total weight of dry fly ash, FGD ash, or mixture thereof so treated.

Claims

1. A method of reducing sulfur concentration in fly ash, flue gas desulfurization (FGD) ash, and mixtures thereof, the method comprising contacting fly ash, FGD ash, or mixtures thereof with an aqueous acidic solution, for a time, at a temperature, and at a liquid-to-solid ratio wherein sulfur concentration within the fly ash, FGD ash, or mixture thereof is reduced to no more than 5 wt % SO.sub.3 based on total weight of dry fly ash, FGD ash, or mixture thereof so contacted.

2. The method of claim 1, comprising contacting the fly ash, FGD ash, or mixtures thereof with the aqueous, acidic solution at a liquid-to-solid ratio of no more than 2 liters aqueous, acidic solution per gram of fly ash, FGD ash, or mixture thereof.

3. The method of claim 1, comprising contacting the fly ash, FGD ash, or mixtures thereof with the aqueous, acidic solution at a liquid-to-solid ratio of no more than 1 liter aqueous, acidic solution per gram of fly ash, FGD ash, or mixture thereof.

4. The method of claim 1, comprising contacting the fly ash, FGD ash, or mixtures thereof with the aqueous, acidic solution at a liquid-to-solid ratio of no more than 100 mL aqueous, acidic solution per gram of fly ash, FGD ash, or mixture thereof.

5. The method of claim 1, comprising contacting the fly ash, FGD ash, or mixtures thereof with the aqueous, acidic solution at a liquid-to-solid ratio of no more than 50 mL aqueous, acidic solution per gram of fly ash, FGD ash, or mixture thereof.

6. The method of claim 1, comprising contacting the fly ash, FGD ash, or mixtures thereof with the aqueous, acidic solution at a liquid-to-solid ratio of no more than 25 mL aqueous, acidic solution per gram of fly ash, FGD ash, or mixture thereof.

7. The method claim 1, wherein pH of the aqueous, acidic solution is from about 2.0 to about 6.5.

8. The method claim 1, wherein pH of the aqueous, acidic solution is from about 4.0 to about 6.0.

9. The method claim 1, wherein pH of the aqueous, acidic solution is from about 4.0 to about 5.0.

10. The method of claim 1, wherein the aqueous acidic solution comprises carbonic acid.

11. The method of claim 10, wherein the fly ash, FGD ash, or mixture thereof is contacted with water in the presence of a gas phase comprising a partial pressure of carbon dioxide of from about 0.12 atm to about 10 atm.

12. The method of claim 10, wherein the fly ash, FGD ash, or mixture thereof is contacted with water in the presence of a gas phase comprising a partial pressure of carbon dioxide of from about 1 atm to about 5 atm.

13. The method of claim 1, wherein the aqueous acidic solution comprises a mineral acid.

14. The method of claim 13, wherein the mineral acid is selected from the group consisting of hydrochloric acid (HCl), nitric acid (HNO.sub.3), phosphoric acid (H.sub.3PO.sub.4), sulfuric acid (H.sub.2SO.sub.4), boric acid (H.sub.3BO.sub.3), hydrofluoric acid (HF), hydrobromic acid (HBr), and perchloric acid (HClO.sub.4).

15. The method of claim 1, wherein the fly ash, FGD ash, or mixture thereof is contacted with the aqueous acidic solution at a liquid-to-solid ratio wherein hydrogen ion concentration ([H.sup.+]) in the aqueous acidic solution is at least 1.95 mmol [H.sup.+] per mole of sulfur to be washed.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0019] FIG. 1 is a graph showing annual worldwide production of FGD ash (in millions of metric tons; grey field) and FGD ash utilization (⋄) and fly ash utilization (.circle-solid.) as a percent of the total produced.

[0020] FIG. 2 is an X-ray diffraction pattern of raw FGD ash showing crystalline mineral phases: calcium sulfite hemihydrate (⋄), magnesia (.star-solid.), quartz (▮), tricalcium aluminate (∇), portlandite (), calcite (.diamond-solid.), magnetite (⋄)

[0021] FIG. 3A is a graph showing the concentration of sulfur in the leachate versus dissolution time. FIG. 3B is a graph depicting the corresponding selectivity as the molar ratio of sulfur and calcium in the leachate.

[0022] FIG. 4A is a graph showing the effect of L/S on weight loss and chemical composition due to washing in DI water and 1 atm CO.sub.2; the red data points correspond to the X-axis. FIG. 4B is a graph depicting the effect of acid addition on sulfur removal from the ash at L/S of 25 mL/g. FIG. 4C is a graph depicting the experimentally determined influence of CO.sub.2 pressure on the solubility of sulfur from SDA ash and simulated solubility curves assuming saturation of CaSO.sub.3.0.5 H.sub.2O.

[0023] FIG. 5A is a graph showing the effect of recycle ratio on the net acid required and the wastewater generated for nitric acid and sulfuric acid-based washing solutions. FIG. 5B is a graph showing the effect of recycle ratio on overall gypsum yield as a ratio of washed ash. FIG. 5C is a graph showing the effect of recycle ratio on the amount of gypsum precipitated as a contaminant in the washed ash.

[0024] FIG. 6 is a graph depicting the preliminary estimate for the cost of acid washing of ashes of various SO.sub.3 content to meet the ASTM C618 standard without recycling of washing solution and using sulfuric acid valued at $100/tonne (i.e., “metric ton”=1,000 kg=1 Mg); costs are normalized to the quantity of washed ash.

[0025] FIG. 7A is a block diagram for acid washing according to the present method using mineral acid. FIG. 7B is a block diagram for two-stage sulfur washing and carbon dioxide capture according to the present method.

DETAILED DESCRIPTION

Abbreviations and Definitions

[0026] ASTM=The former American Society for Testing and Materials, now known as ASTM International, West Conshohocken, Pa., USA.

[0027] DI=de-ionized.

[0028] FGD=flue gas desulfurization.

[0029] ICIS=Independent Commodity Intelligence Services, New York City, N.Y. USA; www.icis.com).

[0030] ICP=inductively coupled spectroscopy. ICP-OES=inductively coupled plasma optical emission spectrometry.

[0031] LOI=loss on ignition.

[0032] L/S=liquid/solid ratio.

[0033] MATS=U.S. Mercury and Air Toxics Standards, 40 CFR Part 63.

[0034] Mineral acid=any acid derived from one or more inorganic compounds. A non-exclusive list of mineral acids that can be used in the present process include (but are not limited to) hydrochloric acid (HCl), nitric acid (HNO.sub.3), phosphoric acid (H.sub.3PO.sub.4), sulfuric acid (H.sub.2SO.sub.4), boric acid (H.sub.3BO.sub.3), hydrofluoric acid (HF), hydrobromic acid (HBr), perchloric acid (HClO.sub.4), and the like.

[0035] Hydroiodic acid HI

[0036] SCM=supplementary cementitious material.

[0037] SDA=spray dryer absorber.

[0038] TGA=thermogravimetric analysis.

[0039] Tonne=metric ton=1,000 kg=1 Mg.

Materials and Methods:

[0040] Sulfur-rich FGD ash was generously provided by Weston Generating Station, Marathon County, Wisconsin, United States. The ash was generated from a semi-dry desulfurization technology, spray dryer absorber (SDA), which was commissioned upstream to pulse-jet filter. This arrangement thus isolates and collects a mixed residue of fly ash and FGD products. The raw ash was characterized using X-ray fluorescence (XRF) spectroscopy (for elemental composition) and X-ray diffraction (XRD) for mineral phase makeup.

[0041] Four sets of dissolution experiments were conducted:

[0042] 1) Batch dissolution of the ash in DI water;

[0043] 2) Batch dissolution of the ash in dilute aqueous nitric acid;

[0044] 3) Titration of ash slurry using aqueous nitric acid; and

[0045] 4) Batch dissolution under 1 to 5 atm of partial pressure of CO.sub.2.

[0046] A typical batch dissolution experiment involved charging a predetermined quantity of FGD ash (measured to the accuracy of 0.001 g) into a 250 mL glass bottle containing DI water (conductivity <1 μS/cm and measured to the accuracy of 0.5%) and stirred at 400±20 rpm on a hot-plate magnetic stirrer. The liquid-to-solid (L/S) ratio was varied in the range of from 25 to 2000 mL/g. The reaction temperature was maintained at 23±2° C. Slurry pH was measured using an Orion Ross™-brand glass electrode and a Thermo Scientific Orion Star pH meter (both obtained commercially from ThermoFisher Scientific, Waltham, Mass., USA) that was regularly calibrated with pH 4 and pH 10 buffers (slope within 99-100%). Each experiment was repeated at least twice, with at least one experiment without the pH probe to avoid KCl contamination due to the electrode and to measure potassium concentration in the leachate. Experiments under carbon dioxide environment (1-5 atm) were conducted in a 50 mL benchtop Parr reactor (Parr Instrument Company, Moline, Ill., USA). For each batch experiment, 25 mL of deionized water was first added to the reactor and charged with ash to achieve a slurry with L/S of 25 mL/g to 100 mL/g. The slurry was stirred for 30 s to avoid agglomeration of the ash. Following this, without stirring, the air in the reactor was evacuated and replaced with 100% CO.sub.2 (×3). The reactor was then finally charged with CO.sub.2 and maintained at the desired CO.sub.2 pressure. Subsequently, dissolution was started by stirring the slurry at 800±10 rpm. The liquid sample was collected at the end of reaction, after depressurizing the reactor.

[0047] Acid titration of the ash slurry was carried out using a Hanna 901C auto-titrator (Hanna Instruments, Smithfield, R.I., USA). with standardized 1.0 M nitric acid as the titrant, and the initial L/S ratio was 25 mL/g. A linear dosing titration was carried out with 0.05 mL every 5 seconds until the slurry pH was reduced to 3.0. Intermittent liquid sampling was carried out for analysis.

[0048] During each batch experiment and titration study, the liquid sample was collected by filtering the slurry using a 0.2 μm syringe filter. An in-sample oxidation procedure of leachate was necessary to oxidize sulfite to sulfate and consequently avoid the formation of sulfur dioxide upon acidification to a pH≤2 for ICP analysis. Without oxidation, a positive error in concentration measurement with significant variance was noticed, possibly due to higher nebulization of volatile dissolved gases into the ICP chamber (Sarudi et al., 2001). The in-sample oxidation procedure involved adding 0.03 wt. % H.sub.2O.sub.2 in 1 mM HNO.sub.3 solution to collected liquid samples in a 1:1 volume ratio. Partial acidification was necessary for instantaneous oxidation. Subsequently, the oxidized sample was diluted and acidified using 0.5 M HNO.sub.3 to match the ICP standards. Elemental concentrations (Ca, Fe, Mg, Al, Si, S, Na, and K) in the leachate were measured using ICP-OES after calibrating with certified standards procured commercially from Millipore-Sigma (St. Louis, Mo., USA) and High-Purity Standards (North Charleston, S.C., USA). The leached residues were collected from reaction slurry after vacuum filtration on a quantitative (1 μm retention) filter paper and dried overnight. Loss on ignition (LOI) was measured as weight loss of ash from thermogravimetric analysis (TGA) in zero-air environment at 1000° C.

[0049] A geochemical model was built using the PHREEQC v3 program (Parkhurst and Appelo, 2013) with the wateq4f thermodynamic database. (As of Apr. 15, 2020, the PHREEQC program and all supporting documentation can be downloaded free of charge from the United States Geological Survey at www.usgs.gov/software/phreeqc-version-3.) The model was used to determine the solubility controlling mineral phases and the transport of elements due to washing. Model results are compared with the experimental data. Stoichiometric acid moles and water requirements determined from the titration and batch dissolution experiments were used to simulate washing requirements to meet the ASTM C618 standard (5% SO.sub.3) under the CO.sub.2 environment and compared with experimental results.

Ash Characterization:

[0050] The elemental composition of the raw ash is shown in Table 1. The mineral phase composition of the raw ash is shown in FIG. 2. FIG. 2 is an X-ray diffraction pattern of the raw FGD ash showing crystalline mineral phases: calcium sulfite hemihydrate (⋄), magnesia (.star-solid.), quartz (▮), tricalcium aluminate (∇), Portlandite (), calcite (.diamond-solid.), and magnetite (⋄). The analyses confirm high-sulfur content (10.44 wt. % as SO.sub.3) in the form of calcium sulfite hemihydrate (CaSO.sub.3.0.5 H.sub.2O), which makes it non-compliant with the ASTM C618 standard (maximum permissible sulfur content is 5.0% as SO.sub.3 wt. % per the standard). Calcium content as CaO in the ash was ˜26 wt. %, which makes it suitable for a Class C SCM. Of the total calcium in the ash, it is estimated that ˜30% of it is associated with calcium sulfite hemihydrate. Other Ca-rich minerals include tricalcium aluminate along with traces of calcium carbonate and portlandite.

TABLE-US-00001 TABLE 1 Chemical composition (wt. %) of ash sample determined using X-ray fluorescence (XRF) Oxide CaO SiO.sub.2 Al.sub.2O.sub.3 SO.sub.3 Fe.sub.2O.sub.3 MgO TiO.sub.2 Na.sub.2O P.sub.2O.sub.5 K.sub.2O BaO SrO Total FGD 25.9 31.8 15.6 10.44 4.94 3.95 1.08 0.97 0.75 0.53 0.45 0.24 96.64 ASH

[0051] It is generally agreed that the sulfur-rich mineral phases are accumulated on the surface of the fly ash, and the grain size of these phases are expected to be smaller than those phases associated with fly ash (Enders, 1996; Izquierdo and Querol, 2012). The sulfur leaching characteristics determined by testing disclosed herein leads to a process scheme to reduce the sulfur content of the ash (and optionally to capture CO.sub.2 at the same time).

Sulfur Extraction Kinetics:

[0052] Dissolution experiments were performed in DI water to understand the rate of sulfur release from the ash and its selectivity. FIG. 3A is a graph showing the concentration of sulfur in the leachate as a function of dissolution time. FIG. 3B is a graphing showing the corresponding selectivity for sulfur as a molar ratio of sulfur and calcium in leachate. Both graphs used DI as the liquid phase. Here, selectivity is defined as the molar ratio of sulfur to calcium in the leachate, which demonstrates the relative extents of dissolution of the calcium sulfite hemihydrate and tricalcium aluminate phases. These are the two main phases as shown in FIG. 2 that contain calcium. The extraction kinetics suggests a rapid initial release followed by a plateau in concentration vs. time. Also, the relative release of sulfur was found to decrease with time, suggesting a relative increase in the rate of tricalcium dissolution with time. These results are due to higher sulfur content on the surface ash, which preferentially leaches from the surface of the ash until it reaches its saturation concentration in the liquid phase. The relative extraction of tricalcium aluminate is lower due to it being embedded within the matrix of the ash particle. In view of retaining tricalcium aluminate in the leached residue, a relatively short dissolution time, five minutes, was used for further studies of selective sulfur removal in a batch dissolution experiment.

Effect of L/S Ratio on Sulfur Extraction Characteristics in DI Water:

[0053] A series of experiments were run using a five-minute dissolution time and using various L/S ratios of the DI water and FGD ash. The leachates were then subjected to elemental analysis (ICP, ICP-OES) to determine the concentrations of selected elements within the leachate. The concentration of various elements (in mg/dm.sup.3=g/m.sup.3) in the leachate released within the first five minutes of the batch dissolution experiments at various L/S ratios is shown in Table 2. As shown in the table, Ca and S are the main elements in the leachate, both when DI is used as the liquid phase and at all the CO.sub.2 pressures tested.

[0054] From FIGS. 3A, 3B, and Table 2, where the effect of L/S on selectivity and leachate concentrations are shown, it is clear that the sulfite phase dissolution selectivity improves with a decrease in leachate pH, which in turn was a consequence of increasing L/S ratio. Thus, it appears that high L/S and acidic conditions favor the selective washing process. Also, in DI water the concentration of Al in the leachate was found to decrease at lower leachate pH, especially in the near-neutral pH range, where the solubility of aluminum hydroxide is the lowest. The observed selectivity behavior can be attributed to enhanced solubility of calcium sulfite hemihydrate vis-à-vis tricalcium aluminate leaching kinetics with decreasing pH and complete solubility of trace mineral phases such as calcite and portlandite at low L/S. This is corroborated by the estimated supersaturation with respect to calcium sulfite hemihydrate, as shown by the saturation indices in Table 2. The calcium sulfite hemihydrate phase rapidly dissolves to saturate the solution, and subsequent leaching of other calcium-rich phases supersaturates it.

TABLE-US-00002 TABLE 2 Effect of L/S (mL/g) on concentration (mg/dm.sup.3) of major elements released from FGD ASH into the leachate at 23 ± 2° C. L/S Ca Si Mg Al Fe Na K S pH.sup.† SI.sub.Hh Deionized water 25 219.22 0.66 0.27 41.32 n.d. 6.29 3.82 94.29 11.24 1.31 200 81.94 1.21 0.58 9.35 n.d. 2.49 0.24 40.25 10.91 0.67 500 54.05 0.82 0.39 4.53 n.d. 2.24 0.25 25.06 10.70 0.33 1333.3 35.66 0.54 0.19 1.98 n.d. 1.89 0.17 19.03 10.43 0.08 2000 28.00 0.48 0.15 1.39 n.d. 1.84 0.16 15.12 10.35 −0.1 100% CO.sub.2 - 1 atm abs. pressure 25 683.1 18.8 35.1 0.4 n.d. 13.8 11.4 252.0 6.15 1.15 50 501.3 12.6 31.9 0.6 n.d. 9.3 10.5 199.6 6.03 0.86 100 431.5 9.7 65.5 6.1 n.d. 12.5 10.3 248.0 5.99 0.86 100% CO.sub.2 - 2.5 atm abs. pressure 25 771.8 20.2 34.4 1.4 n.d. 21.1 20.4 287.8 5.81 0.94 100% CO.sub.2 - 5 atm abs. pressure 25 924.7 26.9 44.5 2.5 n.d. 21.8 20.7 395.7 5.52 0.89 n.d. - not detected/concentration lower than detection limit (0.1 mg/dm.sup.3) .sup.†pH for samples under the CO.sub.2 environment was estimated using PHREEQC assuming CO.sub.2 saturation and are not experimentally measured values. SI.sub.Hh is the saturation index defined as the logarithm of the ratio of ionic activity product and solubility product and estimated using the PHREEQC program.

[0055] The residual mass and composition of the leached ash was estimated by mass balance. The weight of the residual ash after treating with DI water vs L/S ratio are shown in FIG. 4A; the weight percent of sulfur removed with added acid are shown in FIG. 4B, left Y-axis, with selectivity to sulfur shown on the right Y-axis; the results for solubility of CaSO.sub.3.0.5 H.sub.2O as a function of CO.sub.2 pressure are shown in FIG. 4C. The water requirement for reducing the SO.sub.3 concentration below 5.0 wt. % in the washed ash is estimated to be 1300 mL/g of ash. The corresponding weight loss was estimated to be about 15 wt. %. The method thus enables using 85 wt. % of the landfilled ash, which otherwise would be landfilled due to its high sulfur content. As shown in Table 3, the washed ash composition is compliant with the ASTM C618 standard.

[0056] While this outcome is promising, a huge volume of water is required per unit mass of ash. This makes the process using water alone difficult from a regulatory standpoint (if not a cost standpoint). Regulatory agencies such as the U.S. Environmental Protection Agency have strenuous permitting requirements for any industrial process that uses large volumes of water. Therefore, methods were tested to enhance sulfur solubility (selectively) by acidifying the liquid phase.

Sulfur Extraction Characteristics in Acid Media:

[0057] Acid titration using 1.0 N HNO.sub.3 was carried out on ash slurry at initial L/S ratio of 25 mL/g to determine the acid requirement and resulting sulfur selectivity. FIG. 4B shows the experimental extent of sulfur removal and the corresponding sulfur selectivity. Until the point of complete sulfur removal, the sulfur selectivity increases with acid addition. Thus, the data shown in FIG. 4B demonstrate that the calcium sulfite hemihydrate phase is preferentially/selectively dissolved under acidic conditions. The acid requirement for washing to achieve 5.0 wt. % SO.sub.3 in the residue at L/S of 25 mL/g was determined to be at least 1.425 mmol/g of ash; the corresponding slurry pH was 4.64. The stoichiometric requirement of acid is estimated to be 1.95 meq/mmole of sulfur released.

Recycling of Washing Solution:

[0058] To further reduce the water and acid losses, optimal recycling of washing solution was investigated through process simulation studies. Simulations were performed for two washing solutions, nitric and sulfuric acids, where the anion in the latter solution reacts with leached calcium to precipitate sparingly soluble gypsum. In simulation studies, recycle ratio—defined as the ratio of recycle flow rate to overall volumetric flow rate of washing solution into the ash washing unit—was varied to study its effect on various process performance indicators such as net acid required, net water required/wastewater generated, overall gypsum production, and extent of gypsum contamination in washed ash. The results are shown in FIG. 5. The findings from FIG. 5A suggest that, at L/S of 25 mL/g, 98% and 100% recycle is optimal for nitric acid and sulfuric acid washing solutions, respectively; the corresponding wastewater generation rates are 0.5 mL/g and 0 mL/g, respectively. The net acid requirement for nitric acid and sulfuric acid washing solutions are 1.035 and 1.007 meq/mmol of S(IV) washed, respectively, which is about 50% lower than that consumed without any recycling. The gypsum yield is higher in nitric acid-based washing solution compared to that of sulfuric acid. As recycle ratio approaches 1, the gypsum yield is comparable for both acids (see FIG. 5B). Sulfuric acid-based washing solution is promising but for higher gypsum in the washed ash. At 100% recycle of sulfuric acid washing solution, the gypsum contamination is ˜13 mg/g of ash (see FIG. 5C). The contribution of SO.sub.3 content in the washed ash from Gypsum is estimated to be about 0.9 wt % (as SO.sub.3), which still meets the 5.0 wt % upper limit set for SCM in ASTM C618.

[0059] At a competitive price of $90-100 per tonne of sulfuric acid (ICIS, 2018), the putative cost of acid for washing FGD ASH is roughly US $7.10 to $7.90 per tonne of washed ash without recycling washing solution (US $3.70 to $4.10 per tonne of washed ash with optimal recycling of washing solution). See FIG. 6. For FIG. 6, sulfuric acid was priced at $100/tonne without recycling of washing solution. The resulting costs are normalized to the quantity of washed ash. The resulting de-sulfurized ash sells wholesale for about $100 per metric tonne for use in making concrete (which typically costs 17% less than Portland cement) (American Road and Transportation Builders Association, 2011).

TABLE-US-00003 TABLE 3 Chemical composition (in wt. %) of raw ash and washed ashes and compliance with ASTM C618 standard 100% S Acid Sat. CO.sub.2 ASTM removal.sup.‡ DI Water washing washing C618 Element/phase Raw Ash estimate Experimental Standard CaO 25.9 22.8 22.6 21.7 23.1 ≥18.0 SO.sub.3 10.44 0 4.3 5.0 4.9 ≤5.0 SiO.sub.2 + 52.3 64.1 60.3 60.7 60.2 ≥50.0 Al.sub.2O.sub.3 + Fe.sub.2O.sub.3 LOI — — — — — ≤6.0 .sup.‡For an elementary estimate, it is assumed that all S is present as CaSO.sub.3, and its leaching is 100% selective.

[0060] Preliminary estimates, shown in FIG. 6, demonstrate that significant value addition is possible for moderate sulfur ashes (below 20 wt. % SO.sub.3) by acid washing. Oxidation of sulfite to sulfate and subsequent precipitation of gypsum by evaporation or spray drying will attract additional cost.

[0061] A first version of the process based on mineral acid washing is shown in FIG. 7A. A second version of the process based on CO.sub.2/carbonic acid washing is shown in FIG. 7B. Although mineral acid washing appears attractive, in view of identifying a recyclable acidic source to reduce the input cost further, ash washing using an aqueous solution saturated with CO.sub.2 was explored. Such a process allows pure carbon dioxide capture from a dilute flue gas stream and gypsum production from the leachate. First, the effect of L/S (25-100 mL/g) and pressure (1-5 atm) under a 100% CO.sub.2 atmosphere were studied. As shown in FIG. 4A, the effect of L/S has similar characteristics on residual slag in both pure DI water and CO.sub.2 environment. The same extents of sulfur washing in both approaches resulted in similar chemical composition and weight loss. To achieve 5.0 wt. % residual sulfur under a 1 atm CO.sub.2 environment, the L/S ratio was estimated to be about 100 mL/g ash. This is a substantially smaller water requirement (˜13 times smaller than DI water washing per unit mass of ash). As shown in Table 2, the concentration of sulfur in the leachate is invariant when L/S changed from 25 to 100 mL/g under 1 atm CO.sub.2 pressure. These results suggest that the release of sulfur into the solution is limited by solubility. To further improve the sulfur solubility, the influence of CO.sub.2 pressure on the release of sulfur was studied. The results are shown in FIG. 4C. An increase in the CO.sub.2 pressure from 1 atm to 5 atm increased the sulfur solubility by about 57%. Under these conditions, the water requirement was vastly reduced, to approximately 50 mL/g ash.

Process Scheme for Sulfur-Washing Using Carbon Dioxide:

[0062] The block diagram for the proposed combined processes of sulfur washing and carbon dioxide capture is illustrated in FIG. 7B. The first stage involves absorbing carbon dioxide from flue gas into FGD ash slurry to partially dissolve the sulfite phase. In the following stage, sulfur washing is carried out using pressurized carbon dioxide (up to 5 atm) to overcome the low solubility of calcium sulfite hemihydrate in aqueous solution. The low-sulfur residual fly ash is filtered at this stage. The sulfur-rich aqueous stream may optionally be oxidized by peroxide addition, which acidifies the solution and leads to spontaneous desorption of carbon dioxide. The desorbed carbon dioxide is partially compressed and recycled to the sulfur-washing step. The rest is available for CO.sub.2 sequestration or utilization. To reduce peroxide consumption, a limited amount of oxygen gas can be provided. The oxygen supplied should be less than the stoichiometric requirement, to avoid contaminating the CO.sub.2 to be recycled. The choice of using a combination of O.sub.2 gas and peroxide is principally economic, based on the cost of O.sub.2 and peroxide.

[0063] The sulfate solution, which is also rich in calcium, may be concentrated either by evaporating or recycling for subsequent precipitation as gypsum; if required, calcium hydroxide may optionally be added at this stage to maximize gypsum precipitation. The process water recovered is recycled to the CO.sub.2 absorber in the first step or sent to wastewater treatment. The pH of the oxidized liquid stream is expected to be in the range of 3-4. Recycling this stream back into the process will lower the water requirement as compared to a single pass system.

CO.sub.2 Capture Capacity from Flue Gas:

[0064] CO.sub.2 capture from flue gas, which was described as the first stage in the proposed process description, is predicted based on the absorbed CO.sub.2 and alkalinity of the leachate, defined as the molar equivalents of bicarbonate ion concentration in the leachate. While the contribution of physically absorbed CO.sub.2 is estimated using Henry's law, the estimation of alkalinity requires the experimental knowledge of all the cations and anions in the solution. A conservative estimate of alkalinity can be obtained based on the solubility of CaSO.sub.3.0.5 H.sub.2O phase and neglecting the calcium release from other phases. As shown in the below balanced chemical reaction, the concentration of bicarbonate ions in the solution would be equal to that of sulfite ion concentration, which can be estimated using solubility curve for CaSO.sub.3.0.5 H.sub.2O.

CaSO.sub.30.5H.sub.2O+H.sub.2CO.sub.3.fwdarw.Ca.sup.+2+HSO.sub.3.sup.−+HCO.sub.3.sup.−+0.5H.sub.2O

[0065] As shown in FIG. 4C, the solubility data predicted based on thermodynamic data for pure CaSO.sub.3.0.5 H.sub.2O phase underestimates the experimentally determined sulfur concentration. An empirical fit with experimental data can be obtained when the solubility curve is shifted by a constant value to match the experimentally determined sulfur solubility in DI water (P.sub.CO2=0 atm). Based on the empirical fit for the solubility curve, the sulfur release into aqueous solution saturated with the flue gas (12 vol % CO.sub.2) at a L/S ratio of 50 mL/g ash is estimated to be 3.25 mM; corresponding absorbed CO.sub.2, and bicarbonate ion concentrations are 4.07 and 3.25 mM, respectively. The corresponding CO.sub.2 capture potential is estimated to be 16.1 kg/tonne of ash. Because the total dissolved inorganic carbon is only 7.32 mM, any significant CO.sub.2 capture per tonne of ash is possible only at higher L/S values.

Critical Analysis of Water Intensity:

[0066] A preliminary analysis was carried out to examine the water requirement for the proposed process vis-à-vis wet FGD technology deployed in a 500 MW supercritical power plant. This comparison is to understand the water intensity of the process disclosed herein with a known benchmark (wet FGD process). The design basis for wet FGD technology is based on the work by the U.S. Department of Energy (DOE) researchers for a typical supercritical powerplant (Klett et al., 2007). For comparison, the same design parameters given in the DOE study for the ash washing process were used, although the sulfur content in the DOE study is 44.8 wt. %, which is ˜400% higher than the ash used in this study.

TABLE-US-00004 TABLE 4 Comparison of water intensity for Wet FGD and SDA followed by sulfur washing Process parameter Units Value A. Power plant design basis Net power generation MW 500 Net Plant efficiency % 39.9 Coal tonne/day 3877.5 S in coal wt. % of coal 4.35 Fly ash produced wt. % of coal 8.2 B. FGD Wet FGD SDA + W1.sup.‡ SDA + W2.sup.‡ Fly ash tonne/day 316.8 1139.sup.$ 1139.sup.$ % SO.sub.3 in ash* wt. % — 44.7 44.7 Dry gypsum produced tonne/day 1097 0 0 Water Intensity L/MWhr 220 140 140 C. Washing Process — W1 W2 Liquid-to-solid ratio mL/g — 437 437 CO.sub.2 capture tonne/day — 159.1 0 Water losses tonne/day — 426 318 Water intensity L/MWhr — 35.5 26.5 Gypsum produced tonne/day — 1059 1059 D. Overall Marketable ASTM C618 compliant ash tonne/day 316.8 344.6 344.6 Water Intensity L/MWhr 220 175.5 166.5 .sup.‡W1, W2 are two- and single-stage washing scenarios, respectively. *SO.sub.3 content in the ash is much higher compared to the ash used in this study. While the washing process may not be ideal for washing such high sulfur ashes, we intend to estimate and compare water intensity with the wet FGD process. .sup.$Sulfur is assumed to be present as calcium sulfite hemihydrate.

[0067] Two configurations were explored for the sulfur-washing process shown in FIG. 7B: i) W1—both stage 1 and stage 2 for carbon dioxide capture and washing as described in FIG. 7B, and ii) W2—only stage 2 for sulfur washing without carbon dioxide capture. A comparison of process parameters for wet FGD and sulfur washing processes are shown in Table 4. The washing process produces similar quantities of gypsum as wet FGD in addition to CO.sub.2 capture of 159 tonnes/day. The water losses in wet FGD are found to be higher than the two-stage washing process, W1, which in turn is higher than the single-stage process without CO.sub.2 capture, W2. Such a decreasing trend in water intensity is due to lower flue gas flow rates (˜4.6% of total generated) into the absorption column compared to the wet FGD process where entire flue gas stream is contacted with process water. In the case without CO.sub.2 capture, W2, water losses are not expected during absorption stage. A marginal contribution in lowering water intensity for washing processes is due to retention of 5.0 wt. % SO.sub.3 in the washed ash.

Conclusion

[0068] The sulfur extraction from SDA FGD ash is rapid in both alkaline and acidic conditions, as observed in DI water and carbonic acid, respectively. The release of sulfur appears to be limited by the solubility of calcium sulfite hemihydrate. In general, the leachate was observed to be supersaturated with respect to CaSO.sub.3.0.5 H.sub.2O phase, possibly due to the continued release of calcium from other phases and the common-ion effect. The selectivity of sulfur removal is better at lower residence time in the washing reactor, whereby calcium losses can be avoided. The lower solubility of aluminum, iron, and silicon in the acidic conditions reduces the washing losses in sulfuric acid or carbonic acid and shows potential for recovery of gypsum by oxidation of the leachate. The stoichiometric ratio of the acid requirement for S washing was found to be about 1.95 and 1.01 meq/mmole of sulfur released at L/S ratio of 25 mL/g ash without and with recycle, respectively. The net wastewater generation for mineral acid based processes is estimated to be less than 0.5 mL/g of ash. The liquid-to-solid ratio requirement for the FGD ash studied is estimated to be 50 mL/g ash at 5 atm CO.sub.2 pressure. The washed ash is shown to be compliant with ASTM C618 for class C SCM's. The gravimetric losses due to washing are estimated to be 15 wt. % of the FGD ash. Thus, up to 85 wt. % of the ash stored in ponds can be beneficially used as SCM. The process is useful for reducing the sulfur content of sulfur-rich coal ashes from flue gas desulfurization by washing them using mineral or carbonic acids. Because the process equipment required is similar to the equipment used in wet FGD and incinerator ash washing processes, the process can be commercialized on a large scale with only minimal capital investment.

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