Oxy-pressurized fluidized bed combustion: Configuration and options analysis

Applied Energy 262 (2020) 114531

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Oxy-pressurized fluidized bed combustion: Configuration and options analysis

T

Robert T. Symonds , Robin W. Hughes, Margarita De Las Obras Loscertales ⁎

Natural Resources Canada, CanmetENERGY, 1 Haanel Drive, Ottawa, ON, Canada

HIGHLIGHTS

to biomass resulted in reductions in both power and oxygen requirements. • Shifting the operating pressure led to a higher amount of available heat. • Increasing optimum operating pressure between 7 and 15 bar(g) was obtained. • An use of a WSA process for sulphur removal resulted in increases in useable heat. • The • Changes in flue gas O concentration via WSA impacts process performance. 2

ARTICLE INFO

ABSTRACT

Keywords: Industrial CCS Oxy-fuel combustion Pressurized fluidized bed Biomass Wet sulphuric acid process

Bio-energy carbon capture and storage is a promising option to mitigate or eliminate CO2 emissions from fossil fuel-based heat, steam, and power generation. Among the different carbon capture technologies, oxy-fuel fluidized bed combustion is an attractive option due to its fuel flexibility and moderate operating temperature, making it suitable for bio-energy carbon capture and storage. A recent breakthrough in the development of this technology is operation under pressurized conditions. In order to assess the benefits of oxy-pressurized fluidized bed combustion, a number of design configurations and options were analyzed via process simulation. Three coals, one coke, and one torrefied hardwood biomass were selected for this work. A series of parameters, including reactor and system configuration, were varied to determine their impact on key performance metrics such as oxygen requirement, power balance, availability of high temperature heat, and by-product generation. The shift from coal to torrefied biomass only marginally affected the total amount of recoverable heat for power production, but resulted in significant reductions in both power and oxygen requirements. It was observed that an increase in operating pressure led to a higher amount of available heat due to the recovery of the latent heat of H2O condensation. An optimum operating pressure, ranging from 7 to 15 bar(g), was obtained when considering the total power consumed within the power plant. Results showed that the total quantity of recoverable high temperature heat is slightly higher for the circulating bed configuration in comparison to that of the bubbling. The use of a flue gas wet sulphuric acid process for sulphur removal resulted in the production of a high purity sulphuric acid product (> 98 wt%) and increases in useable heat were noted at elevated pressures where total sulphur capture could be pushed to as high as 99.97%. Additionally, changes in flue gas O2 concentration via the wet sulphuric acid process and sorbent injection into the combustor were found to have significant impacts on process performance and natural gas/O2 demands.

Abbreviations: ASU, air separation unit; BECCS, bio-energy carbon capture and storage; CAPEX, capital expenditures; CCS, carbon capture and storage; CFBC, circulating fluidized bed combustion; CPF, central processing facility; CPU, CO2 processing unit; DCC, direct contact cooler; FBC, fluidized bed combustor; FEHE, feed effluent heat exchanger; FG, flue gas; EHX, external heat exchanger; GHG, greenhouse gas; HOR, heat of reaction; In-bed HX, in-bed heat exchanger; NGCC, natural gas combined cycle; OCAC, oxygen carrier assisted combustion; OPEX, operating expenditures; Oxy-FBC, oxy-fluidized bed combustion; Oxy-PFBC, oxy-pressurized fluidized bed combustion; OTSG, once-through steam generator; RFG, recycled flue gas; SAGD, steam assisted gravity drainage; WSA, wet sulphuric acid ⁎ Corresponding author. E-mail address: [email protected] (R.T. Symonds). https://doi.org/10.1016/j.apenergy.2020.114531 Received 28 October 2019; Received in revised form 8 January 2020; Accepted 12 January 2020 0306-2619/ Crown Copyright © 2020 Published by Elsevier Ltd. All rights reserved.

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1. Introduction

and premature failure if not suitably managed [37,38]. By operating at elevated pressures, additional options become available that turn the sulphur-based components from waste (e.g., CaSO4 from calcium-based sorbent inject [39]) to valuable by-products. One such option is the oxidation of SO2 to SO3 and the subsequent hydration of SO3 to form sulphuric acid – the so called wet sulphuric acid (WSA) process [40,41]. The objective of this work was to assess the potential benefits of utilizing biomass-based and high sulphur containing feedstocks in the context of oxy-FBC. While these fuels have been considered in the past, as noted above, significant knowledge gaps exist on their respective performance under pressurized oxy-fuel conditions, including the variety of additional downstream processing options that become available at elevated pressures. In this paper, we evaluated a series of potential feedstocks and oxy-PFBC configurations to determine their suitability for low GHG emitting heat, steam, and power production. Included in this work are design specifications, configuration options, process simulation methodology, and case study outputs used to examine the impact of key operating parameters and flue gas processing options on high temperature heat availability and power consumption within the plant. In addition, a comparison between bubbling and circulating fluidized bed configurations is made that evaluates their respective characteristics versus those expected from an ideal oxy-fuel combustion system.

Worldwide awareness of the importance for greenhouse gas (GHG) reduction from fossil fuel-based heat, steam, and power generation has significantly increased in recent years. There has been a substantial R& D effort in developing carbon capture and storage (CCS) technologies, but CCS technologies using bio-energy sources have received far less attention [1]. Biomass represents a promising alternative to fossil fuels for energy applications with the potential to reduce the atmospheric CO2 concentration, achieving an overall negative carbon balance on a life cycle basis [2,3], when CCS is employed [4]. Bio-energy accounted for approximately 10% of the total final energy consumption and 1.4% of global power generation in 2015 [5]. Studies reported by McIlveenWright et al. [6] showed that co-firing coal with biomass in circulating fluidized bed combustion (CFBC) power plants could lower CO2 emissions in comparison to coal alone and meet EU legislation on NOX and SOX emissions with only a slight decrease in electricity generation efficiency. Several authors have noted that the use of biomass presents some challenges associated with lower energy densities [7], fouling [8,9], agglomeration [10,11,12,13] and corrosion [14,15]. However, the potential for generating negative CO2 emissions with very low SOX and NOX production makes it a very attractive option for heat, steam, and power production provided energy efficiency penalties could be mitigated. There are several approaches to improve biomass combustion performance, but they are all typically related to material densification and/or pre-treatment. In the case of densification, the biomass bulk density is increased to improve downstream steps such as storage, handling, and injection [16]. Biomass pre-treatment options include: (1) washing to remove adverse ash components [17], (2) torrefaction to reduce moisture content and increase heating value [18,19], and (3) steam explosion to alter the physical/chemical structure resulting in a high density and low moisture reabsorbing end-product [20]. Each approach varies in terms of energy consumption, resulting in a trade-off between level of biomass enhancement and marginal gain in global process efficiency. Therefore, process performance of pre-treated biomass must be compared to traditional fuel feedstocks to assess its true value and commercial viability. Oxy-fluidized bed combustion (oxy-FBC) is a CO2 capture technology that has been successfully demonstrated at the 30 MWth scale using coal [21,22] and is one of the most developed CCS technologies [23]. It is now considered both technically feasible and economically competitive for commercial applications [24,25]. Unlike conventional air-fired combustors, oxy-fuel combustors utilize purified oxygen instead of air as oxidant to avoid dilution (via N2) of the flue gas. After water and impurity removal high-purity CO2 is produced, ready for compression and sequestration/utilization [26,27]. Since fluidized bed combustion has several advantages over other combustion methods, including fuel flexibility and moderate operating temperature, it is expected that this technology is suitable for bio-energy CCS (BECCS) [28]. Recent techno-economic analyses have shown that operating oxyFBC systems at elevated pressures is a relatively attractive option from both cost and efficiency perspectives [29]. This is in large part because this suite of technologies can successfully integrate the heat of condensation of the water vapour present in the flue gas into the power cycle [30,31,32] or into boiler feed water preheat [33]. In the past, this heat was often considered waste heat [34], but through pressurization the heat is available for use at a higher temperature. It stands to reason then that pressurized oxy-combustion may also be an attractive option for BECCS given the relatively high water vapour content of biomassbased flue gas. In addition to biomass-based feedstocks, more challenging fossil fuels could also be promising options for oxy-pressurized FBC (oxyPFBC), specifically high sulphur containing ones such as bituminous coals, petroleum cokes, and hydrocarbon residues [35,36]. The high sulphur content of these fuels can lead to equipment corrosion, fouling,

1.1. Design specification For each case assessed in this work, the total solid fuel thermal input to the combustor was fixed at 100 MWth. This value was selected, as it is in line with the typical heat, steam, and/or power demands of several different industrial processes. Although not an exhaustive list, some of the applications currently being considered include: (1) replacement of once-through steam generators (OTSGs) at steam assisted gravity drainage (SAGD) heavy oil extraction facilities, (2) combined heat and power for oil sands partial upgraders, (3) heat generation at pulp and paper mills, and (4) various other processes that require small- to midrange power generation. It should be noted that larger-scale power generation remains an option, but within the Canadian context, it is unlikely that oxy-PFBC processes could compete with natural gas combined cycles (NGCCs) [42] that currently meet the Government of Canada’s GHG emission-intensity limit (420 t CO2/GWh) without any post-combustion CO2 capture [43]. Since many of these industrial applications require a far lower thermal input than typical commercial-scale power generating stations, this allows for the construction of significantly smaller combustors – the size is further reduced by operating at elevated pressures. This result is beneficial given the remote locations of many of these facilities, where location cost factors and labour costs are relatively high. By limiting the maximum size of all pressure vessels and rotating equipment, shop assembly and transport by road or rail becomes possible, thereby reducing costs. Although not discussed in detail in this work, curbing the total capital expenditures (CAPEX) of oxy-fuel based technologies is likely required before commercial uptake is realized [44]. 1.2. Oxy-PFBC configurations Internationally, there are two broad classes of fluidized beds under evaluation for pressurized oxy-fluidized bed combustion: bubbling and circulating. The long-term objective of this study is to provide a side-byside comparison of the two approaches to determine which of these technologies are best suited to various fuels and applications based on the calculated cost of heat, steam, and/or power. This paper presents the first phase of the study in which configurations and operating conditions are evaluated for future economic analysis. The pressurized oxy-FBC configurations considered in this study are presented in Fig. 1 (bubbling) and Fig. 2 (circulating). The key differences between the two processes are related to heat extraction and sulphur removal strategies. 2

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Fig. 1. Base case oxy-bubbling PFBC process flow diagram; ASU - air separation unit, In-bed HX – in-bed heat exchanger, Conv HX - convective heat exchanger, Cond HX - condensing heat exchanger, DeOxo - catalytic deoxidation.

The circulating FBC utilizes an external heat exchanger (EHX) for the majority of the heat extraction (via an inert circulating bed material), whereas this is done via an in-bed heat exchanger (In-bed HX) in the bubbling FBC case. When comparing the two, the underlying assumption was that potential changes to the combustor and process configuration should improve the overall economics of oxy-FBC, while ensuring that system reliability is maintained, if not improved. To help guide this development, the following key design requirements (Table 1), along with expected characteristics, have been selected to represent the ideal fluidized bed-based oxy-fuel combustion system: Several of the expected characteristics can be at least partially met by operating at elevated pressures. The reduced volumetric flow rate minimizes the combustor diameter and the overall dimensions of much of the downstream stream equipment (e.g., cyclones, direct contact cooler (DCC), etc.). In addition, heat transfer rates are known to be enhanced at elevated pressures [45,46,47] which facilitates a more compact boiler design. The recovery of higher temperature latent heat, which is only possible in pressurized systems, can also improve that overall process efficiency. While these process benefits are, for the most

part, configuration independent, many of the remaining expected characteristics are dependent on the sulphur removal strategy. Since the bubbling bed configuration employs heat transfer surfaces within the dense bed region, in-situ sulphur capture is essentially the only option to protect against corrosion. It should be noted that in-situ capture is also an option for CFBs. The most common approach is to inject a calcium-based sorbent, such as limestone or dolomite, into the bed to capture the majority of the sulphur [48,49,50]. The sulphation reactions are as follows: Direct: CaCO3 + ½O2 + SO2 → CaSO4 + CO2 mol CaCO3 Indirect: CaCO3 ↔ CaO + CO2 CaO + ½O2 + SO2 → CaSO4

ΔHro = −323.1 kJ/ (1)

ΔHro = +178.0 kJ/mol CaCO3 ΔHro

(2)

= −501.1 kJ/mol CaO (3)

Sulphation using CaCO3 can proceed via two paths where the difference lies in the intermediate formation of CaO through calcination (Equation (2)). CaCO3 calcination is an equilibrium constrained

Fig. 2. Base case oxy-circulating PFBC process flow diagram; ASU - air separation unit, EHX - external heat exchanger, IHX - internal heat exchanger, Conv HX convective heat exchanger, Cond HX - condensing heat exchanger, DeOxo - catalytic deoxidation. 3

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Table 1 Design requirements and expected characteristics of an ideal fluidized bed-based oxy-fuel combustion system. Design Requirements Compact boiler Fuel conversion Scalable Temperature control

Expected Characteristics

heat transfer area (increased overall heat transfer coefficients) • Minimized of fluidizing velocity and firing rate • Decoupling > 99.5% consistently observed • Conversion is a function of equipment design • Smaller industrial applications (e.g., SAGD and pulp & paper) to large-scale power generating stations (< 100 MW ) • Uniform temperature distribution and heat transfer ( ± 2 °C) • Low NO (keep combustor temperature below ~950 °C) • Oxidant staging (multiple injection ports along the reactor height) • Low risk of agglomeration (keep combustor temperature below ash fusion temperature) • High velocity (> 2 m/s) • Smaller footprint/lower cost • Limited heat transfer surfaces in dense bed regions • Sulphur capture (> 95%) • Low sulphur containing fluidizing gas (high efficiency sulphur removal before recycling) • Limited exposure to SO /SO and high moisture containing flue gas (reduced risk of acid condensation) • As low as 20–30% of full firing rate • Low ash fuels (< 10 wt%, ideally < 2 wt%) • Reduced sorbent requirements (switch to other methods of sulphur capture) • High temperature heat generation (> 250 °C) • Low sensible heating demand (< 5% of total heat available) • th

X

Minimized combustor diameter Minimized erosion Minimized heat exchanger corrosion Load following Minimized solids handling High efficiency

2

3

2NO2 + 2NaOH ↔ NaNO3 + NaNO2 + H2O ΔHro = −163.1 kJ/mol NO2 (6)

reaction and is a function of both temperature and CO2 partial pressure [51]. In most cases, the conditions expected during oxy-pressurized combustion will preclude the possibility of indirect sulphation due to both high combustion temperatures (typically > 850 °C) and CO2 partial pressures. However, this is not necessarily the case for atmospheric (or low pressure) oxy-combustion and could lead to lower system efficiencies. In these cases, extra heat is required for the endothermic calcination of CaCO3 to CaO, but large quantities of CaO are not utilized for sulphur capture. Past work on pressurized CFBs have shown that Ca/S ratios of > 3 are required to attain more than 95% sulphur retention [52,53]. This lowers the amount of high temperature heat available in the combustor. It should also be noted that sorbent injection increases sensible heating demands in the combustor, increases the total quantity of CO2 requiring purification and compression, increases the total volumetric flow of flue gas (i.e., increased vessel dimensions), and adds to the solids handling requirements. When in-situ capture is not a requirement, several additional sulphur removal techniques are possible, such as caustic addition and acid production through the WSA process. Caustic addition is likely the simplest method and takes advantage of the high-pressure conditions [54,55]. A process flow diagram of this sulphur removal approach is shown in Fig. 3. For low sulphur containing fuels, such as biomass, or in cases where in-situ sulphur capture is sufficiently high, a condensing heat exchanger (Cond HX1) can be placed upstream of the scrubber to recover latent heat for boiler feed water (BFW) pre-heating (or a suitable alternative). Restricting the SOX concentration entering the condensing heat exchanger is critical in mitigating the risk of acid condensate corrosion [56]. While this risk could be moderated with enhanced materials, the increase in costs would likely be prohibitive. The flue gas exiting the condensing heat exchange is sent into a scrubber system (direct contact cooler – Cond HX2) with a recirculating water stream to recover the sensible and remaining latent heat from near the pinch point down to below the water dew point temperature. The scrubbing system can effectively remove fine particulate matter, which may damage downstream rotating equipment, and can be used for both SOX and NOX removal to meet CO2 purity specifications [57]; Kinder [58] via caustic solution addition (e.g., NaOH): SO2 + 2NaOH ↔ 2Na2SO3 + H2O ΔHro = −1337.9 kJ/mol SO2 SO3 + 2NaOH ↔ 2Na2SO4 + H2O ΔHro = −1811.4 kJ/mol SO2

NO2 + NO + 2NaOH ↔ 2NaNO2 + H2O ΔHro = −274.7 kJ/mol NO (7) The water condensed from the flue gas and that generated by SOX/ NOX capture (Eqs. (4)–(7)) can be directed to a steam cycle or central processing facility (CPF) in the case of SAGD applications. This is an important result as it reduces the fresh water make-up requirements of the facility and improves the overall environmental sustainability of the process. Unfortunately, while the scrubbing system is very effective, and required for final flue gas polishing, one cannot take advantage of the highly exothermic nature of the sulphur capture reactions from a process efficiency perspective (i.e., all low-grade waste heat). Acid production through the WSA process is a potential sulphur removal technique that can be used in conjunction with a scrubbing system to meet the design specifications of pressurized oxy-fuel combustion systems. Fig. 4 provides a process flow diagram for the WSA process. As a first step, the flue gas enters a SO2 converter unit where it is catalytically converted to SO3 and subsequently H2SO4 via the following reactions: SO2 + 1/2O2 ↔ SO3

ΔHro = −99.1 kJ/mol SO2

(8)

SO3 + H2O → H2SO4

ΔHro

(9)

= −133.0 kJ/mol SO3

The SO3-rich flue gas leaving the converter is sent to a second unit (WSA Condenser) where it is subsequently hydrated (Eq. (9)) to form sulphuric acid (H2SO4), condensed, concentrated, pumped, and finally cooled resulting in a high purity commercial-grade product (> 95 wt% H2SO4 (aq)). Several advantages are realized by adopting this technique for SOX removal in addition to producing a valuable product. The first is that this process can handle flue gas streams with high sulphur concentrations (> 10,000 mg SO2/Nm3) [59]. This opens the door to the utilization of low cost, high heating value feedstocks, such as petroleum cokes and hydrocarbon residues, without the efficiency penalties associated with sorbent injection. Another advantage of the WSA process is that all required reactants (O2 and H2O) are already present in sufficient quantities in the flue gas leaving the combustor. Therefore, there is no need for additional injection equipment and no losses to sensible heating. Furthermore, the SO2 converter makes use of the excess O2 in the flue gas, which reduces the amount of natural gas required for catalytic deoxidation (DeOxo) without affecting combustion performance. The final advantage is related to the production of high

(4) (5) 4

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Fig. 3. Flue gas condensation process flow diagram for moisture, SOX/NOX, and particulate matter removal; CPF - central processing facility, BFW – boiler feed water.

temperature heat. By operating the SO2 converter above 300 °C, it is possible to benefit from the exothermic sulphur capture reactions. This heat can be used for BFW pre-heat or recycled flue gas (RFG) re-heat, for example. Higher temperature operation is possible (400–500 °C), but at the cost of poorer SO2 conversion. Fortunately, operating at elevated pressures favours the production of SO3 and therefore helps mitigate this adverse effect. The increase in operating pressure also provides a benefit to the WSA condenser by increasing the acid dew

point temperature. Laursen [59] calculated the latent heat of condensation of H2SO4 to be ~69 kJ/mol, which is now available for integration within a steam cycle or CPF. Material selection for the condenser is an important consideration, as it will likely have the largest impact on total costs for the WSA process. Rosenberg [40] has suggested the use of acid and shock resistant boron-silicate glass for the condenser tubes with an acid resistant brick-lined collection drum, so costs should be manageable even at elevated pressures.

Fig. 4. Wet sulphuric acid (WSA) process flow diagram; CPF - central processing facility, RFG - recycle flue gas.

5

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The flue gas exiting the scrubber system should meet pipeline specifications for SOX and NOX, but the oxygen concentration would almost certainly be in the percentage level and; therefore, would be well above the acceptable limit of 10 (Type I and II pipelines) to 50 ppm (Type III pipelines) [60,57]. There are several different methods for oxygen removal, but the most promising option appears to be via a catalytic deoxidation reactor (DeOxo, see Figs. 1 and 2) [61,62]. In this process, the residual oxygen in the flue gas is reacted with a gaseous fuel in the presence of a catalyst. Although H2 has been used in this process at the commercial scale, natural gas is likely a more appropriate option given the remote locations of many of the potential oxy-FBC applications, coupled with the cost of producing high purity H2 (i.e., impurities in the H2 could impact CO2 quality and compression performance). Eq. (10) shows the reaction for methane-based flue gas deoxidation. It should be noted that natural gas contains other higher hydrocarbons (C2-C5s) that are also expected to act as deoxidation reagents. CH4 + 2O2 → CO2 + 2H2O ΔHro = −890.3 kJ/mol CH4

Table 2 Power correction factors for oxygen production at elevated pressures [65]. Oxygen Production Pressure (barg)

Power Correction Factor

1 5 7 10 15 20 30 40

1.0285 1.0808 1.1050 1.1390 1.1896 1.2334 1.3034 1.3547

2. Process simulation methodology 2.1. Fluid package and thermodynamic data Pressurized oxy-FBC simulations were performed with Aspen HYSYS V9 using the Peng-Robinson fluid package. Where available, process modeling parameters were selected using a process quality guideline for energy systems document prepared by the U.S. National Energy Technology Lab [66], including recommended efficiencies for turbomachinery (only adiabatic efficiencies considered), approach temperatures for heat exchange equipment, and pressure drops. A pressure drop of 50 kPa was assumed across the combustion equipment including losses from gas distribution, combustor bed/freeboard, and solids separation cyclones. Additionally, no heat loss to the surroundings was considered. As previously discussed, limestone was utilized for sulphur capture (sorbent injection cases only) and an olivine sand (Mg2SiO4) used as the circulating bed material for circulating bed configuration cases. To ensure accurate predictions of heats of reaction and solids sensible heating/cooling, heat capacity and heats of formation data was obtained from FactSage 7.3. This was required, as Aspen HYSYS does not include suitable thermodynamic databases for the solid compounds. Thermodynamic data for the species used in this work can be found in Appendix A.

(10)

The optimal approach to heat management in the deoxidation system is largely conditional on two factors: (1) the inlet conditions of the flue gas and (2) the process-wide heat integration strategy. The first factor is the most critical as it dictates the total amount of heat available and the maximum temperature of this heat. The outlet temperature of the scrubber system is somewhat flexible, but temperatures in the range 50–90 °C are required if reasonable water knock-out is to be achieved (< 4 vol% H2O). This means that for relatively low excess oxygen conditions, the maximum deoxidation temperature would be limited (< 300 °C) and would likely favour an adiabatic type reactor arrangement. On the other hand, when more heat via deoxidation is available the preferred option would likely be an isothermal reactor with a feed effluent heat exchanger (FEHE). This high temperature heat (> 300 °C) could be directly utilized in steam cycles or more advanced cycles such as those incorporating supercritical CO2 [61]. Although not explicitly shown in the process flow diagrams (Figs. 1 and 2), a CO2 compression train is required to bring the CO2 product to pipeline pressure and remove the remaining moisture (down to < 30 ppmv). Integrally geared compressors can be used, as they have been shown to minimize operating and capital costs for CO2 compression over two-casing single-shaft compressors. After bulk moisture removal in a flash vessel, the CO2 product requires further drying to meet the specification; power requirements for drying are estimated to be 0.000136 kWeh/mol of feed gas [63]. The final major consideration in any oxy-fuel combustion configuration is the oxidant source. There are several different approaches to oxygen generation including cryogenic distillation, membrane separation, pressure-swing adsorption, vacuum-swing adsorption, molten salts, etc. While each approach offers different advantages, oxygen production via cryogenic air separation units (ASUs) is likely the most applicable to industrial applications due to its commercial availability, high purity, and high volume production rates [64]. This not only reduces oxy-fuel technology risks, but also ensures CO2 products specifications can be met (< 4 vol% N2 for Type I and II pipelines). Unfortunately, oxygen production is the single largest power consumer in any oxy-fuel system and, therefore, has a major impact on the overall system efficiency. Zheng [65] estimates the power consumption of ASU oxygen production at approximately 230 kW·h/tonne of O2 at atmospheric pressure, where Table 2 can be used to correct for elevated pressure operation. It is therefore critical to consider configuration options that can minimize oxygen demands, without adversely affecting process performance or reliability. In addition, integrating waste heat from the ASU and using the nitrogen by-product for fuel drying (as shown in Figs. 1 and 2) are potential options for decreasing the energy penalty associated with oxygen production.

2.2. Reactants and base case definition Several solid fuel feedstocks covering a wide range of higher heating values (HHV), oxygen, and sulphur content were evaluated. Specifications for each fuel are provided in Table 3. All solid fuels and limestone entered the process at 20 °C and 1.01 bar(a) and were raised to combustor pressure via lock hoppers before injection. High pressure O2 was generated using a cryogenic ASU. The oxygen flow rate was varied to maintain 1.0–5.0 vol% O2 (2.0 vol% O2 base case) in the flue gas leaving the combustor and the recycled flue gas flow rate was adjusted to maintain inlet O2 concentrations ranging from 10 to 50 vol% (40 vol% base case) entering the combustor. A convective heat exchanger (Conv HX) was used to lower the flue gas temperature to 350 °C before entering a condensing heat exchanger (Cond HX1) and direct contact cooler (Cond HX2) to recover the heat of condensation of the water vapour and drop the flue gas water content to 3.7 vol%. The condensing heat exchanger was replaced with a SO2 converter and acid condenser in cases where WSA was utilized for sulphur removal. The operating temperatures of the SO2 converter was maintained at 400 °C, while the outlet temperature of the acid condenser was varied to meet a > 98 wt% H2SO4 acid product stream. The heat generated during the WSA process was used for recycle gas pre-heating. The O2 remaining in the flue gas was eliminated via a catalytic deoxygenation (DeOxo) reactor using pure CH4 as the fuel resulting in the formation of CO2 and H2O. It is expected that the natural gas supply to any commercial oxy-FBC facility would contain some level of N2, CO2, and higher hydrocarbons. However, these species have been omitted as they do not strongly influence process performance and do not significantly change the flue gas composition entering the 6

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Table 3 Solid fuel specifications (as received). Type

Torrefied Biomass Hardwood

Boundary Dam Coal Lignite

Highvale Coal Sub-bit

Illinois #6 Coal Bituminous

Delayed Coke Petroleum

Proximate (wt%) Moisture Ash Volatile Fixed Carbon

3.68 1.12 74.03 21.17

4.21 14.31 39.32 42.16

4.18 22.86 29.46 43.50

5.90 8.88 38.40 46.82

5.96 1.06 8.52 84.47

Ultimate (wt%) Carbon Hydrogen Nitrogen Sulphur Oxygen (by difference) Mass HHV (MJ/kg)

51.68 5.94 0.13 0.01 38.56 20.26

59.86 3.81 1.02 0.91 15.88 23.23

53.64 3.30 0.72 0.31 14.99 20.56

68.55 4.70 1.40 3.11 7.46 28.90

83.44 3.35 1.71 4.45 0.04 33.80

compression train. When required, the O2-free flue gas exiting the DeOxo reactor was subsequently compressed (using inter-stage cooling) to 25 bar(g), dried, and then further compressed to meet pipeline specifications [57]. Details pertaining to O2 and natural gas conditions and compositions at the battery limits of the process are provided in Appendix B.

elevated pressure operation [65]. 3. Results and discussion 3.1. Oxy-bubbling PFBC One of the key benefits of pressurized oxy-fuel combustion over that of non-pressurized technologies is the potential for increased high temperature heat recovery. Fig. 5 shows hot composite curves at various operating pressures for the torrefied biomass described in Table 3. As expected, increasing the operating pressure increases the amount of recoverable heat for power production [68]. The majority of this increase is due to the recovery of latent heat of H2O condensation that increases high temperature heat recovery from ~86 to 94% when increasing the operating pressure from 1 to 40 bar(g). It is worth noting the significant increase in high temperature recoverable heat in the combustor itself (i.e., heat available at 900 °C). This result, while beneficial from a system efficiency perspective, will have combustor design implications. Higher operating pressures will reduce the size of the combustor and will limit the maximum heat transfer surface area available for exchange within the riser. If this is the case, operating the combustor as a CFB, as opposed to a bubbling fluidized bed (BFB), may be prudent even if it reduces the temperature at which heat can be recovered, as is necessary with external heat exchange in the CFB configuration. Although an optimum operating pressure cannot be determined solely from the results depicted in Fig. 5, increasing the pressure from 30 to 40 bar(g) did not result in a substantial change in recoverable heat. This is the result of H2O phase equilibria; the steam saturation temperature does not change greatly with increasing pressure above 30 bar(g) and suggests pressures beyond this value should not be considered. Details pertaining to material battery limits, energy balances, and power balances for all cases can be found in Appendix C. To further the discussion on impact of operating pressure on oxyPFBC performance, it is important to consider how net power production changes. Fig. 6 plots the power consumed within the power plant versus operating pressure for the torrefied biomass. It was determined that only the oxygen production (via ASU) and the flue gas compression/drying consumed significant power. The recycle blower power is reduced when increasing the operating pressure, but its contribution is less than 5% of the total power consumed by the power plant even at 1 bar(g). It should be noted that placement of the recycle blower upstream and downstream of the DCC was considered. It was determined that placing the recycle blower after the DCC provided greater benefits including reducing acid dew point risk (i.e., corrosion versus materials of construction) and reducing the actual volumetric flow of flue gas via flue gas moisture removal and temperature reduction. From Fig. 6, it can be seen that the optimum operating pressure appears to fall between 7 and 15 bar(g), which matches well with previous studies [30,69,32]. At pressures below 7 bar(g), the power requirements for

2.3. Case summary In order to assess the suitability of oxy-PFBC for a variety of applications, a series of parameters, including reactor and system configuration, were varied to determine their impact on key performance metrics such as oxygen requirement, power balance, availability of high temperature heat, and by-product generation. These parameters included: 1. 2. 3. 4. 5. 6.

Operating pressure Fuel type Inlet oxygen concentration Outlet oxygen concentration Reactor configuration strategy Sulphur removal strategy

Table 4 summarizes the process simulation case definitions for all parameters considered. The combustor temperature was fixed at 900 °C with full fuel conversion for all cases. While SOX formation and its impact on process performance is discussed in detail, NOX formation is not considered. Instead, the fuel nitrogen for the various fuels simply appears as N2 in the combustor flue gas. The actual level of NOX in the flue gas is expected to be at a level low enough not to affect process performance and would be removed in the DCC along with any remaining SOX. Cases 1–16 and 18 consider sorbent injection as the primary sulphur removal strategy, when required. Here, limestone (pure CaCO3) feed rates were adjusted to maintain a SO2 concentration of 100 ppmv in the flue gas leaving the combustor. A 40% limestone utilization was assumed [53] in all cases. Olivine sand was used with circulating fluidized bed configurations (Cases 16–24) and its circulation rate was adjusted to maintain a net zero heat duty in the bed portion of the combustor and 700 °C exiting the EHX. Typical olivine circulation rates ranged between 260 and 280 kg/s and 1% of the recycled flue gas was directed to the EHX as fluidizing gas. In Cases 1–6 the combustor operating pressure was varied from atmospheric pressure to 40 bar(g) (15 bar(g) base case). The utilization of a DCC for primary sulphur removal (Cases 20–24) was only considered with low sulphur containing torrefied biomass; the remaining high sulphur cases (17 and 19) employed WSA. In this study, the assumed pinch point was 130 °C, which is in line with a typical steam Rankine cycle [67] and SAGD applications [33]. Power demands for oxygen production and flue gas compression/drying were determined via correlation and corrected for 7

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5 4 3 1 2 2

CFB DCC CFB DCC CFB DCC CFB DCC CFB DCC

2

BFB Sorb. Inject. 2

BFB Sorb. Inject. 2

CFB Sorb. Inject. 2

CFB WSA

CFB Sorb. Inject. 2

CFB WSA

flue gas processing outweigh that of oxygen production. The opposite is true when the pressure is increased beyond 15 bar(g), albeit the difference is less dramatic. Therefore, at higher pressures, heat integration with a steam power cycle, or supercritical CO2 in more advanced cycles, now becomes ever more important [70]. The increase in high temperature heat available for power production at higher pressure could offset the increased power demand, provided the additional heat can be effectively integrated. Process economics will play an important role in determining the optimal operating pressure, specifically the capital costs associated with pressurized O2 production. Previous studies have indicated potential issues related to a shift from coal-based to biomass-based heat/power production, most notably a concern with power density. Torrefaction of biomass has the potential to overcome this concern [71], provided system operation and performance is not excessively altered. When comparing hot composite curves between coals and torrefied biomass, the total amount of recoverable heat for power/steam production (assuming a 130 °C pinch point) only dropped from 90.3% in the case of Illinois #6 to 89.7% for torrefied biomass (Fig. 7). In addition, the torrefied biomass has a nearly identical performance to both Highvale and Boundary Dam coals in this regard. Only small differences in total recoverable heat in both the combustor and convective heat exchanger are observed when comparing the torrefied biomass to these coals. However, Illinois #6 and the delayed coke perform slightly better for two reasons. The first is related to the substantially higher heating values (see Table 3), which minimizes the sensible heating requirement to bring the fuel up to combustion temperature. It should be noted that this is somewhat offset by the higher oxygen mass flowrates required for complete combustion. The second reason is that both Illinois #6 and the delayed coke have large sulphur contents leading to substantial sorbent injection requirements. This does increase the sensible heating requirements in the combustor, but it is more than made up for by the heat generated via direct sulphation (Eq. (1)). Although not explicitly explored as one of the cases in this work, it has been observed that operating at lower pressure conditions gives the opposite result. Here, sorbent heating and calcination demands (Eq. (2)) exceed the heat generated via sulphation. This further justifies operating oxy-fuel systems at elevated pressures, as improving sorbent utilization is likely not feasible [53]. As was the case when considering changes in system performance by varying the operating pressure, changes in facility power consumption is a useful metric when comparing differences between fuel types. Fig. 8 depicts both the power consumed and oxygen required versus fuel type for the oxy-bubbling PFBC process at 15 bar(g). Again, only the oxygen production and flue gas compression/drying are significant contributors to the total power consumed. Although the flue gas compositions differ between the five fuels, the quantity of power consumed in flue gas compressions/drying are very similar ( ± 3%) and only account for ~25% of the total power demand. This suggests that oxygen production is ultimately controlling the difference in power demand. As shown in Fig. 8, the oxygen required to achieve full fuel conversion is significantly higher for all coals/coke studied in comparison to that of the torrefied biomass, even when accounting for the lower heating value of the biomass. This is primarily attributed to the lower oxygen content of the coals/coke; 38.6 wt% versus 7.5 wt% for torrefied biomass and Illinois #6 coal, respectively, and lower still for the delayed coke. Illinois #6 performed marginally better than Highvale and Boundary Dam coals in this respect because of its combination of high carbon and low ash content. As an example, if one were to switch from firing Boundary Dam coal to torrefied biomass, a ~0.68 MWe savings could be achieved, which would equate to an approximate 2% increase in total power generation efficiency (HHV basis). This result is quite significant and does not include other cost saving benefits of torrefied biomass. As discussed above, the sulphur content of torrefied biomass is relatively low in comparison to the other coals and coke considered. For

Outlet O2 Concentration (% wet)

TB = Torrefied Biomass, HC = Highvale Coal, BDC = Boundary Dam Coal, I6C = Illinois #6 Coal, DC = Delayed Coke.

BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2 BFB Sorb. Inject. 2

BFB Sorb. Inject. 2

TB 40 TB 40 TB 40 TB 40 TB 40 BDC 40 BDC 40 DC 40 DC 40 TB 50 TB 30 TB 21 TB 16 TB 10 DC 40 I6C 40 BDC 40 HC 40 TB 40 TB 40 TB 40 TB 40 TB 40

TB 40

15 15 15 15 15 15 15 15 15 15 15 15 15 15 15 15 15 15 40 30 20 1 15

Operating Pressure (bar(g)) Fuel Type Inlet O2 Concentration (% wet) Reactor Configuration Sulphur Removal

7

9 8 7 6 5 4 3 2 1 Case

Table 4 Process simulation case definition (variations from base case are underlined and marked in bold).

10

11

12

13

14

15

16

17

18

19

20

21

22

23

24

R.T. Symonds, et al.

8

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1000 900 800

Temperature (oC)

700 600 500 400

1 bar(g)

7 bar(g)

15 bar(g)

20 bar(g)

30 bar(g)

40 bar(g)

300 200 100 0 0

10

20

30

40

50

60

70

80

90

100

110

Heat Load (MWth) Fig. 5. Hot composite curves versus operating pressure – Oxy-bubbling PFBC with torrefied biomass – Cases 1–6.

testing [71]. Thus far, only a relatively high inlet O2 concentration (40 vol%) has been considered. Fig. 9 plots the hot composite curves versus inlet O2 concentration over a wide range of possible inlet O2 conditions at 15 bar(g). From this figure, two important observations can be made, namely: (1) there appears to be a minimum inlet O2 concentration requirement and (2) there is minimal gain in available high temperature heat above ~40 vol% O2. There has been some discussion on whether or not high inlet O2 operation could lead to agglomeration in the dense bed region of the combustor [28]. For this reason, O2 concentrations lower than that of air have been proposed as a method to control particle temperature (i.e., dilution of the oxidizing gas with recycled flue gas). When evaluating this option, it was noted that dropping the

example, delayed coke produces approximately 250 times more SO2 than the biomass assuming full fuel conversion. Since the predicted flue gas sulphur concentrations for torrefied biomass are so low, direct sorbent (Ca-based) injection into the fluidized bed can be completely avoided. This (1) reduces capital costs and process complexity associated with solids injection systems, (2) reduces DCC water clean-up, and (3) largely mitigates in-bed and condensing heat exchanger acid dew point and corrosion issues. Therefore, from a high-level process perspective, there are major advantages in using torrefied biomass in place of conventional coals/cokes when considering oxy-bubbling PFBC technologies. However, differences in torrefied biomass devolatilization and char combustion kinetics, along with potential agglomeration issues, are still largely unknown and require attention via experimental

Power Consumed within the Facility (MWe)

12.5

12.0

11.5

11.0

10.5

10.0

9.5 0

5

10

15

20

25

30

35

40

Operating Pressure (bar(g)) Fig. 6. Total power consumption versus operating pressure – Oxy-bubbling PFBC with torrefied biomass – Cases 1–6. 9

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1000 900 800

Temperature (oC)

700 600 Torrefied Biomass Highvale Coal Boundary Dam Coal Illinois #6 Coal Delayed Coke

500 400 300 200 100 0 0

10

20

30

40

50

60

70

80

90

100

110

Heat Load (MWth) Fig. 7. Hot composite curves versus fuel type – Oxy-bubbling PFBC at 15 bar(g) – Cases 1 and 7–10.

inlet O2 from 16 vol% to 10 vol% resulted in conditions where maintaining a 900 °C combustor temperature was not possible. The sensible heating requirements (O2 + recycled flue gas) outweigh the heat generated via combustion. Furthermore, even under conditions where the required combustor temperature is met (> ~12 vol% O2), the majority of the heat is likely not available at temperatures high enough for optimal integration into a supercritical steam power cycle. However, lower temperature applications such as bitumen extraction via SAGD and heat generation at pulp and paper mills might result in a better fit. In addition to heat integration concerns, large recycle gas requirements increase overall plant CAPEX and operating expenditures (OPEX)

due to greater recycle blower demands. Increasing the inlet O2 concentration from 10 to 40 vol% drops the flue gas recycle blower power requirement by over 85%. Any further increase in inlet O2 provides a negligible benefit from a power savings perspective. It is somewhat unknown at this time if high inlet O2 concentration fluidized bed combustion will result in substantial particle agglomeration, but preliminary testing at CanmetENERGY’s 50 kWth pilot-plant suggests that 50 vol% inlet O2 concentration operation is feasible [72]. As stated above, there is a marginal bump in available high temperature heat above ~40 vol% O2 and operating beyond this point could lead to other potential issues, aside from agglomeration and defluidization. It is possible that by dropping the flue gas recycle flow further, the resulting

9.0 8.8

10.8

8.6 10.6

8.4 8.2

10.4

8.0 10.2

7.8 Power Consumed

10.0

7.6

Oxygen Required

Oxygen Required (kg/s)

Power Consumed Within the Facility (MWe)

11.0

7.4 9.8

7.2

9.6

7.0 Torrefied Biomass

Highvale Coal

Boundary Dam Illinois #6 Coal Coal

Delayed Coke

Fuel Type Fig. 8. Total power consumption and oxygen requirement versus fuel type – Oxy-bubbling PFBC at 15 bar(g) – Cases 1 and 7–10. 10

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1000 900 800

Temperature (oC)

700 600 500 10 vol.%

400

16 vol.% 21 vol.%

300

30 vol.% 200

40 vol.% 50 vol.%

100 0 0

10

20

30

40

50

60

70

80

90

100

110

Heat Load (MWth) Fig. 9. Hot composite curves versus inlet O2 concentration – Oxy-bubbling PFBC with torrefied biomass at 15 bar(g) – Cases 11–15.

combustor diameter (assuming a bubbling bed regime is maintained) could be reduced to a point where (1) fluidization becomes unstable (slugging, wall effects, poor gas distribution, etc.) and (2) there is not adequate space for in-bed heat exchange equipment. Therefore, any potential CAPEX savings via combustor size reduction would not be realized.

geometries, smaller combustors, and therefore, higher firing rates. Since this arrangement does not rely on heat transfer surfaces in the combustion zone, there is greater flexibility when it comes to sulphur removal strategy, specifically for higher sulphur containing fuels. This potentially eliminates, or at least lessens, the need for direct sorbent injection into the combustor, avoiding costly solids handling equipment (injection and high temperature letdown). It also opens the door to higher sulphur containing feedstocks, such as cokes and hydrocarbon residues, where downstream processing via WSA production actually results in a value-added product (commercial grade sulphuric acid) in place of a waste (ash). Fig. 10 plots the hot composite curves versus process configuration and sulphur removal strategy for both the delayed coke (A) and Boundary Dam coal (B) at 15 bar(g). Here, CFB operation is directly compared to that of BFB operation using sorbent injection for sulphur removal. The CFB configuration with sorbent injection is then compared to WSA. In comparison to the BFB cases, there is a downward shift in temperature of recoverable heat due to the lower temperature operation of the external exchanger. However, the total amount of heat available above 700 °C is actually slightly higher (~80.3 MWth CFB versus ~78.5 MWth BFB for delayed coke and ~75.7 MWth CFB versus ~73.6 MWth BFB for Boundary Dam coal). This is an important result as the absolute quantity of this high temperature heat is directly related to the overall efficiency of the steam power cycle. Unfortunately, it does come at the cost of additional heat exchanger surface area requirements because of the lower temperature gradient. In lower temperature applications, this additional surface area would likely be minimal due to a combination of sufficiently high temperature gradients and enhanced heat transfer via high-pressure operation. This means that, from a purely performance-based perspective, the selection between combustor operating modes is likely application specific. Hence, future work should examine the cold composite curves from these applications versus the hot composite curves using a detailed heat integration approach (e.g. pinch analysis) to determine the optimal operating mode. As stated above, WSA sulphur removal is possible when operating the combustor in CFB mode and the performance results depicted in Fig. 10 are quite promising. There is a further increase in available heat above 700 °C (~81.9 MWth for delayed coke and ~77.6 MWth for

3.2. Oxy-circulating PFBC The previous section considered an oxy-bubbling PFBC system and the impact of several key operating parameters (e.g., pressure, fuel type, and inlet O2 concentration) on two major performance metrics, namely: availability of high temperature heat and power consumption. Although not yet analyzed from a quantitative techno-economic perspective, several key conclusions can be drawn: (1) Operating at elevated pressures substantially increases the availability of high temperature heat (> 5%), primarily due to the recovery of latent heat. (2) Elevated inlet O2 concentrations increase the proportion of usable high temperature heat and reduces power demands caused by sensible heating in the combustor. (3) Biomass, in this case torrefied wood, appears to be a promising substitute for coals/cokes due to its similar power density, high oxygen content, and low sulphur content; the latter resulting in significant advantages from process reliability and costs perspectives. Here, potential changes to the combustor and process configuration as a method to improve the economics and reliability of pressurized oxy-fuel combustion are considered. It was anticipated that by operating the combustor in circulating fluidized bed mode that several of the design requirements in Table 1 could be better met in comparison to bubbling mode. By transferring the primary heat removal equipment to an EHX (see Fig. 2), issues related to temperature control, erosion, and corrosion are largely avoided, while the addition of a cyclone should result in higher carbon conversion (via recirculation of unreacted fuel particles). In addition, the higher superficial velocities lend themselves to more compact boiler 11

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1000 900 800

Temperature (oC)

700 600 500 Delayed Coke - BFB - Sorbent Injection

400

Delayed Coke - CFB - Sorbent Injection 300

Delayed Coke - CFB - WSA

200 100 0 0

10

20

30

40

50

60

70

80

90

100

110

100

110

Heat Load (MWth) 1000 900 800

Temperature (oC)

700 600 500 Boundary Dam Coal - BFB - Sorbent Injection

400

Boundary Dam Coal - CFB - Sorbent Injection

300

Boundary Dam Coal - CFB - WSA 200 100 0 0

10

20

30

40

50

60

70

80

90

Heat Load (MWth) Fig. 10. Hot composite curves versus process configuration and sulphur removal strategy – (A – Top) Oxy-PFBC with delayed coke at 15 bar(g) – Cases 10, 16, and 17 – (B – Bottom) Oxy- PFBC with delayed coke at 15 bar(g) – Cases 8, 18, and 19.

Boundary Dam coal) in comparison to when sorbent injection is utilized, which is attributed primarily to lower sensible heating demands in the combustor. However, this is not the only important observation noted when switching to WSA. The conversion of SO2 to SO3, and subsequent hydration, generates additional heat at an optimal temperature for RFG pre-heating. This largely offsets the losses from not generating high temperature heat via sorbent sulphation and is certainly more attractive than removing sulphur via caustic addition in the DCC which operates between ~60 and 90 °C. Other observations included (1) lower flue gas flowrates, (2) lower oxygen demands, (3) lower natural gas demands, and (4) less waste heat; all of which improve process performance and economics, all else being equal. Since these benefits were notably less significant for the Boundary Dam coal,

due to its significantly lower sulphur content (see Table 3), only results for the delayed coke are highlighted in the following discussion. Decreases in various flow rates and demands mentioned above are directly related to the production of H2SO4, which drops both the oxygen and moisture content of the flue gas. This effectively lowers the excess O2 (and total O2 demand by ~0.5%) reducing the risks associated with reducing zones in the combustor. The lower moisture containing flue gas also reduces the cooling demand in the DCC by ~6%. However, the most significant changes are related to the operation of the DeOxo reactor. The lower oxygen content results in a greater than 13% decrease in natural gas demands, causing a net ~2% decrease in total flue gas flow rate (CO2 product). This drops both the power and cooling requirements for compression without sacrificing 12

WSA Heat Recovery (MWth)

6

300

5

250

4

200

HOR - SO2 Converter

3

150

HOR - WSA Condenser Heat - Acid Condensation

2

100

Acid Condensation Temperature 1

50

0

Acid Condensation Temperature (oC)

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0 0

5

10

15

20

25

30

35

40

Operating Pressure (bar(g)) Fig. 11. WSA heat recovery and acid condensation temperature versus operating pressure – Oxy-circulating PFBC with delayed coke.

CO2 product quality. It should be noted that these results could be further enhanced by utilizing even higher sulphur containing fuels. Thus far, the performance results related to WSA as a potential sulphur removal strategy have only been presented at an operating pressure of 15 bar(g) and, while the results appear to be very promising, it is prudent to consider the implications of operating at both lower and higher pressures. Fig. 11 plots the WSA process heat recovery and acid condensation temperature versus operating pressure from atmospheric pressure to 40 bar(g). At each operating pressure, the process conditions have been optimized to achieve the highest level of SO2 conversion, while meeting the commercial grade sulphuric acid specification (~98 wt% H2SO4). The highest level of available heat is realized at the lowest pressure, but as previously discussed; the acid condensation temperature (~110 °C) is well below that of the pinch temperature (130 °C). As the operating pressure is increased, a minimum is reached at ~15 bar(g) with marginal gains thereafter. This suggests that the ideal WSA pressure is primarily a matter of selecting a high enough condensation temperature suitable for heat integration while balancing the penalties associated with excessively high-pressure operation (see Fig. 6). Interestingly, the heat of reaction (HOR) split between the SO2 converter and the WSA condenser diverges as the operating pressure increases. Since no SO2 oxidation takes place in the condenser, this means that the majority of the SO3 hydration is shifted to the converter unit and the associated heat of reaction from hydration is now available at 400 °C. Not only is this advantageous from a process efficiency perspective, but it also has the potential to simplify the WSA process via the combination of the converter and condenser into one reaction vessel. In addition to the level and quality of heat recovery, the conversion of SO2 to SO3 is favoured at elevated pressures. Gibbs free energy analysis predicts that the level of SO2 conversion can be increased from ~98.0 to 99.97% by raising the SO2 converter operating pressure from atmospheric to 40 bar(g). This lessens, if not eliminates, the need for caustic addition in the DCC typically required for final flue gas polishing. The WSA approach to sulphur removal resulted in significant benefits through the reduction of the flue gas oxygen concentration. Therefore, it was of particular interest to explore how changing the combustor outlet oxygen concentration would influence overall process performance. Fig. 12 shows the hot composite curves versus outlet O2

concentration for the torrefied biomass at 15 bar(g). It should be noted that all cases up until this point considered an outlet O2 concentration of 2 vol% and it remains to be seen if such a condition would result in adequate fuel conversion without excessively long residence times. However, it is expected that the circulating fluidized bed arrangement would outperform that of the bubbling configuration due to recirculation of unburnt fuel particles back into the combustor. It is for this reason, in conjunction with the absence of heat transfer surfaces in the bed region, that low excess O2 conditions are considered here. As expected, dropping the outlet O2 concentration to 1 vol% increased the availability of high temperature heat, reduced the O2 and natural gas demands, and dropped the overall power consumption within the facility, but a careful examination of the material, heat, and power balances (see Appendix C) reveals marginal benefits in these regards. It is therefore likely that the added operational risks and potential for poorer carbon conversion outweigh the slight increase in overall process efficiency. Conversely, increasing the outlet O2 concentration from 2 to 5 vol% has serious process implications; ~4% increase in O2 demands, ~160% increase in natural gas demands, ~11% increase in total waste heat, and a ~1% (nominal) drop in total high temperature heat recovery. These results suggest that reducing the minimum required outlet O2 concentration is a critical parameter in maximizing the overall efficiency of any oxy-fuel combustion system. One potential method for achieving this is via oxygen carrier assisted combustion (OCAC) where the typically inert bed material is replaced with a reactive one such as iron or ilmenite ore [73,74,75,76]. These materials have been shown to reduce emissions of unburnt hydrocarbons, increase agglomeration resistance, and reduce corrosion issues. These benefits were noted to be even more pronounced in the dense bed region under low excess oxygen conditions [72]. 4. Conclusions and recommendations An analysis of pressurized oxy-combustion systems has been performed taking into account the effect of key parameters via process simulation. Throughout a series of process simulation cases, it was observed that an increase in operating pressure led to a higher amount of heat being recoverable for steam and power generation, mainly due to the recovery of the latent heat of H2O condensation. An optimum 13

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1000 900 800

Temperature (oC)

700 600 1 vol.%

500

2 vol.% 400

3 vol.% 4 vol.%

300

5 vol.%

200 100 0 0

10

20

30

40

50

60

70

80

90

100

110

Heat Load (MWth) Fig. 12. Hot composite curves versus outlet O2 concentration – Oxy-circulating PFBC with torrefied biomass at 15 bar(g) – Cases 20–24.

operating pressure, ranging from 7 to 15 bar(g), was obtained when considering the total power consumed within the power plant. For a given pressure, oxygen production was the main factor influencing power consumption within the plant since flue gas compressions/ drying demands were found to be similar for all cases at a given pressure. It was observed that the shift from coal to torrefied biomass only marginally affected the total amount of recoverable heat for power production, but resulted in significant reductions in both power and oxygen requirements. Higher inlet O2 concentration operation led to an increase in available high quality heat and a decrease in total power demands. It is expected that the optimum concentration is around 40 vol% O2 to take advantage of the benefits of reduced recycled flue gas flow and to avoid potential issues related to particle agglomeration. Furthermore, O2 concentrations below ~12 vol% must be precluded from consideration as they result in non-feasible operating conditions. A comparison between system configurations, i.e., bubbling versus circulating, yielded several important observations. Preliminary results showed that the total quantity of recoverable high temperature heat (> 700 °C) is slightly higher in the circulating case in comparison to that of bubbling, but this heat is no longer available at the combustion temperature. By eliminating in-bed heat transfer surfaces, higher sulphur containing fuels could be explored without the need for sorbent injection and the added risk of equipment corrosion. The use of a flue gas wet sulphuric acid process resulted in the production of a high purity sulphuric acid product (> 98 wt%). Increases in useable heat were noted at elevated pressures and total sulphur capture could be pushed to as high as 99.97% when operating at pressures up to 40 bar (g), likely eliminating the need for any further downstream flue gas polishing related to sulphur removal. Furthermore, changes in flue gas O2 concentration via the wet sulphuric acid process and sorbent injection into the combustor were found to have significant impacts on process performance and natural gas/O2 demands.

Future studies should include further comparisons between bubbling and circulating fluidized bed configurations and the integration of heat streams into a variety of industrial applications including power generation via a supercritical steam cycle, steam assisted gravity drainage bitumen extraction, and pulp & paper production. Promising applications can be further developed via techno-economic assessments incorporating detailed reactor and equipment sizing with the ultimate objective of estimating key performance metrics such as cost of power and/or steam. These evaluations should consider all major operating costs including O2 production, water and flue gas treatment, and solids disposal. Credit authorship contribution statement Robert T. Symonds: Conceptualization, Methodology, Formal analysis, Writing - original draft, Writing - review & editing. Robin W. Hughes: Conceptualization, Resources, Supervision, Project administration. Margarita De Las Obras Loscertales: Methodology, Resources. Declaration of Competing Interest The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper. Acknowledgement This research was sponsored and funded by the Program for Energy Research and Development (PERD) at Government of Canada, Natural Resources Canada. The authors would like to thank Guy Veilleux at Airex Énergie for providing the torrefied biomass used in this study.

14

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Appendix A. Thermodynamic data See Tables 5 and 6.

Table 5 Heat capacity parameters for various solids. Species

CaCO3 CaSO4 Mg2SiO4

Common Name

Calcium Carbonate Calcium Sulphate Olivine (Forsterite)

Temperature Range (oC)

25–1227 25–1427 25–1727

Cp = a + b * T + c * T^2 + d * T^3 + e * T^4 (kJ/kg·K) a

b

c

d

e

0.489 0.791 0.309

1.173E−03 −9.746E−04 2.441E−03

2.238E−07 3.367E−06 −2.476E−06

−9.636E−10 −2.526E−09 1.186E−09

3.518E−13 5.824E−13 −2.126E−13

Table 6 Heat of formation parameters for various solids. Species

CaCO3 CaSO4 Mg2SiO4

Temperature Range (oC)

Common Name

Calcium Carbonate Calcium Sulphate Olivine (Forsterite)

ΔH = a + b * T + c * T^2 (kJ/kg·mol)

25–1227 25–1427 25–1727

a

b

c

−1,236,860 −1,466,593 −2,221,764

90.214 94.620 136.891

0.0166 0.0325 0.0169

Appendix B. Reactant data See Tables 7 and 8. Table 7 ASU oxygen specifications at battery limits. Oxidant

O2 from ASU

Temperature (°C) Pressure (kPa(a))

−155.0 110

Mole Fraction N2 Ar O2

0.0200 0.0300 0.9500

Table 8 Natural gas specifications at battery limits. Fuel Gas

Natural Gas

Temperature (°C) Pressure (kPa(a)) Mass Higher Heating Value (kJ/kg)

−0.2 5101 55,150

Mole Fraction CH4

1.00

Appendix C. Material, energy, and power data See Tables 9–11.

15

16

0*

0*

100.0

27,177 169

27,172 168

100.0

0

0

9523 0

0

0

9642 0

0

0

35,129 187

0

0

35,069 187

17,587

2

17,587

1

100.0

9646 0

35,101 187

0*

27,172 168

0

0

0

0

17,587

3

100.0

9666 0

35,062 187

0*

27,179 170

0

0

0

0

17,587

4

100.0

9692 0

35,026 187

0*

27,174 168

0

0

0

0

17,587

5

100.0

9728 0

34,994 187

0*

27,174 168

0

0

0

0

17,587

6

100.0

6249 0

36,584 4443

388

29,287 160

0

0

0

17,511

0

7

100.0

6414 0

36,416 3445

1073

29,525 179

0

0

15,497

0

0

8

100.0

6422 0

33,970 4531

2994

29,264 150

0

12,457

0

0

0

9

* Levels of SOX in the flue gas are sufficiently low to not required sorbent injection. ** Sorbent injection replace with WSA.

Outputs Flue Gas (kg/h) Ash/Limestone (kg/h) Water (kg/h) Sulphuric Acid (kg/h) Total Mass Balance (%)

Inputs Torrefied Wood (kg/h) Highvale Coal (kg/h) Boundary Dam Coal (kg/h) Illinois #6 Coal (kg/h) Delayed Coke (kg/h) Oxygen (kg/h) Natural Gas (kg/h) Limestone (kg/h)

Case

Table 9 Oxy-PFBC process simulation battery limits.

100.0

4191 0

35,393 4317

3675

29,417 149

10,652

0

0

0

0

10

100.0

9581 0

35,021 187

0*

27,101 144

0

0

0

0

17,587

11

100.0

9609 0

35,012 187

0*

27,094 149

0

0

0

0

17,587

12

100.0

9620 0

35,022 187

0*

27,109 153

0

0

0

0

17,587

13

100.0

9633 0

35,042 187

0*

27,138 160

0

0

0

0

17,587

14

100.0

9667 0

35,099 187

0*

27,208 177

0

0

0

0

17,587

15

100.0

4191 0

35,380 4317

3675

29,417 149

10,652

0

0

0

0

16

100.0

3836 1477

100.0

6321 0

36,395 3447

1075

0** 34,674 113

29,442 159

0

0

15,497

0

0

18

29,333 129

10,651

0

0

0

0

17

100.0

6222 440

36,164 2218

0**

29,421 153

0

0

15,497

0

0

19

100.0

9711 0

35,048 183

0*

27,152 163

0

0

0

0

17,587

20

100.0

9455 0

34,813 183

0*

26,814 83

0

0

0

0

17,587

21

100.0

9938 0

35,308 183

0*

27,528 252

0

0

0

0

17,587

22

100.0

10,152 0

35,561 183

0*

27,892 339

0

0

0

0

17,587

23

100.0

10,218 0

35,814 183

0*

28,257 425

0

0

0

0

17,587

24

R.T. Symonds, et al.

Applied Energy 262 (2020) 114531

1

100.00 2.58 2.40 0 0 0

65.72 0 19.07 7.58 5.90 1.83 0 0 5.09 100.2 89.7

13

100.00 2.34 2.48 0 0 0

43.14 0 35.01 12.26 8.06 1.61 0 0 5.05 100.3 87.8

Case

Inputs Solid Fuel (MWth) Gas Fuel (MWth) Compression (MWth) Sulphation (MWth) SO2 Converter (MWth) WSA Condenser (MWth)

Outputs In-bed HX (MWth) EHX (MWth) Conv. HX (MWth) Cond. HX (MWth) DCC (MWth) DeOxo (MWth) Acid Condensation (MWth) Acid Cooling (MWth) Waste Heat (MWth) Total Energy Balance (%) High Temperature Heat Recovery (%)

Case

Inputs Solid Fuel (MWth) Gas Fuel (MWth) Compression (MWth) Sulphation (MWth) SO2 Converter (MWth) WSA Condenser (MWth)

Outputs In-bed HX (MWth) EHX (MWth) Conv. HX (MWth) Cond. HX (MWth) DCC (MWth) DeOxo (MWth) Acid Condensation (MWth) Acid Cooling (MWth) Waste Heat (MWth) Total Energy Balance (%) High Temperature Heat Recovery (%)

Table 10 Oxy-PFBC process simulation energy balances.

17 57.68 0 24.75 8.68 7.30 1.71 0 0 5.07 100.3 88.5

100.00 2.45 2.41 0 0 0

14

64.86 0 18.89 6.48 8.58 1.40 0 0 7.54 100.8 85.7

100.00 2.59 4.30 0 0 0

2

70.42 0 15.75 7.31 4.68 1.94 0 0 2.10 100.1 90.8

100.00 2.70 2.35 0 0 0

15

65.24 0 19.01 6.59 7.45 1.68 0 0 5.70 100.5 88.0

100.00 2.58 2.60 0 0 0

3

0 73.49 18.78 4.23 4.05 1.40 0 0 5.19 101.0 92.3

100.00 2.28 2.52 1.32 0 0

16

65.91 0 19.11 8.96 4.25 1.90 0 0 4.87 100.1 91.4

100.00 2.58 2.35 0 0 0

4

0 75.31 16.43 0 3.81 1.15 3.57 0.25 4.71 100.0 91.6

100.00 1.98 2.48 0 0.41 0.39

17

66.26 0 19.16 10.24 2.69 1.96 0 0 4.51 99.9 93.1

100.00 2.58 2.32 0 0 0

5

0 68.22 20.55 4.47 5.60 1.51 0 0 5.88 100.8 89.9

100.00 2.44 2.55 0.39 0 0

18

66.46 0 19.25 10.86 1.96 2.01 0 0 4.24 99.9 94.0

100.00 2.58 2.34 0 0 0

6

0 70.13 18.54 0 5.53 1.43 4.50 0.08 5.85 101.0 89.9

100.00 2.35 2.53 0 0.22 0.02

19

63.78 0 21.51 6.63 5.57 1.51 0 0 6.53 100.4 88.9

100.00 2.45 2.55 0.14 0 0

7

0 67.78 18.95 5.03 7.16 1.57 0 0 4.85 100.4 89.0

100.00 2.50 2.39 0 0 0

20

66.20 0 20.43 6.61 5.58 1.52 0 0 5.88 100.8 89.9

100.00 2.45 2.55 0.39 0 0

8

0 68.05 18.75 5.07 7.04 0.47 0 0 4.69 100.4 89.1

100.00 1.27 2.37 0 0 0

21

68.19 0 19.56 6.41 5.58 1.41 0 0 5.29 100.6 90.3

100.00 2.29 2.46 1.08 0 0

9

0 66.79 19.69 4.96 7.52 2.79 0 0 5.03 100.5 88.7

100.00 3.87 2.42 0 0 0

22

71.80 0 18.52 6.20 4.02 1.41 0 0 5.19 101.0 92.3

100.00 2.28 2.52 1.32 0 0

10

0 66.13 20.17 4.90 7.78 3.98 0 0 5.20 100.5 88.4

100.00 5.19 2.45 0 0 0

23

−19.23 0 79.05 27.64 11.32 1.49 0 0 5.04 100.3 84.7

100.00 2.21 2.78 0 0 0

11

0 65.43 20.70 4.89 7.98 5.16 0 0 5.38 100.5 88.3

100.00 6.51 2.48 0 0 0

24

26.94 0 46.45 16.25 8.90 1.56 0 0 5.05 100.3 87.0

100.00 2.29 2.56 0 0 0

12

R.T. Symonds, et al.

Applied Energy 262 (2020) 114531

Applied Energy 262 (2020) 114531 10.19 10.16 10.13

Supplementary data to this article can be found online at https:// doi.org/10.1016/j.apenergy.2020.114531.

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4.94

11.99

2.58

10.08

10.18

10.15

10.28

10.39

10.74

10.76

10.60

10.74

10.49

10.27

10.19

10.12

10.06

10.74

10.71

10.76

10.74

10.10

10.08

2.62 2.61 2.60 2.59

0.006 0.62 6.43

3.14

2.39

2.10

1.89

2.65

2.65

2.57

2.62

2.58

2.58

2.58

2.58

2.58

2.62

2.58

2.65

2.63

2.60

0.003 0.08 7.47 0.003 0.07 7.45 0.003 0.07 7.42

0.003 0.08 7.48

Appendix D. Supplementary material

Demand DCC Pump (MWe) FG Recycle Blower (MWe) ASU (Production + Compression) (MWe) CPU (Compression + Drying) (MWe) Total Power Demand (MWe)

0.001 0.07 7.44

0.002 0.13 6.91

0.001 0.05 7.71

< 0.001 0.03 8.15

< 0.001 0.03 8.47

0.002 0.08 8.01

0.002 0.08 8.04

0.002 0.08 7.95

0.001 0.08 8.05

0.002 0.48 7.44

0.001 0.25 7.44

0.001 0.18 7.44

0.001 0.10 7.44

0.001 0.04 7.44

0.001 0.08 8.05

0.001 0.08 8.05

0.001 0.08 8.04

0.002 0.08 8.04

0.003 0.07 7.44

23 22 2 1 Case

Table 11 Oxy-PFBC process simulation power balances.

3

4

5

6

7

8

9

10

11

12

13

14

15

16

17

18

19

20

21

24

R.T. Symonds, et al.

18

Applied Energy 262 (2020) 114531

R.T. Symonds, et al.

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19