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Combustion of solid recovered fuels within the calcium looping process – Experimental demonstration at 1 MWth scale
Journal Pre-proofs Combustion of Solid Recovered Fuels within the Calcium Looping Process Experimental Demonstration at 1 MWth Scale Martin Haaf, Jens Peters, Jochen Hilz, Antonio Unger, Jochen Ströhle, Bernd Epple PII: DOI: Reference:
S0894-1777(19)31925-9 https://doi.org/10.1016/j.expthermflusci.2019.110023 ETF 110023
To appear in:
Experimental Thermal and Fluid Science
Received Date: Revised Date: Accepted Date:
12 November 2019 11 December 2019 11 December 2019
Please cite this article as: M. Haaf, J. Peters, J. Hilz, A. Unger, J. Ströhle, B. Epple, Combustion of Solid Recovered Fuels within the Calcium Looping Process - Experimental Demonstration at 1 MWth Scale, Experimental Thermal and Fluid Science (2019), doi: https://doi.org/10.1016/j.expthermflusci.2019.110023
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Combustion of Solid Recovered Fuels within the Calcium Looping Process - Experimental Demonstration at 1 MWth Scale Martin Haaf1*, Jens Peters1, Jochen Hilz1, Antonio Unger2, Jochen Ströhle1, Bernd Epple1 1Institute
for Energy Systems and Technology, Technische Universität Darmstadt, Otto-BerndtStraße 2, 64287 Darmstadt, Germany
2SUEZ
Deutschland GmbH, Binnenhafenstraße 19, 68159 Mannheim, Germany
*Corresponding Author: Martin Haaf, Institute for Energy Systems and Technology, Technische Universität Darmstadt Otto-Berndt-Straße 2, 64287 Darmstadt, Germany Tel.: +49 6151 / 16 22675 Fax: +49 6151 / 1622690 Email: [email protected] Abstract
The calcium looping (CaL) process is an efficient post-combustion CO2 capture technology based on the reversible carbonation-calcination reaction of natural lime. This work presents the results obtained during long-term operation at the 1 MWth CaL pilot plant at Technische Universität Darmstadt. During more than 230 h of representative CaL operation, the calciner was heated by oxy-fuel combustion of waste derived fuels. During the experimental investigation, two types of solid recovered fuels (SRF) were utilized. Both types of SRF were fed to the process in the form of raw fluff, similar to typical industrial applications. The flue gas to be decarbonized in the carbonator, was supplied by an on-site combustion chamber. The CO2 concentration was kept between 9.5 and 10.5 vol.% as typical value for waste-to-energy (WtE) plants fueled by municipal solid waste (MSW). Over a wide range of operation conditions, CO2 absorption rates of 80 % and total CO2 capture rates over 90 % were achieved. Within this work, exemplary long-term data plots of relevant CaL process parameters are shown for each type of SRF. A chlorine balance for a representative operation period is established based on the relevant effluent streams from the CaL system. Thereby, it was found that the calciner fly ash represents the major chlorine effluent. Furthermore, the retention rate of chlorine was above 82 % throughout all test points, respectively. When taking into account the organic waste fractions typically contained in SRF and MSW, this work successfully demonstrates the feasibility of net negative CO2 emissions by means of the CaL process at semiindustrial scale.
1
Keywords: calcium looping (CaL) process, carbon capture and storage (CCS), pilot plant, waste derived fuel, solid recovered fuels (SRF), negative emission technology (NET)
Symbols Ecalc
calciner efficiency, %
Ecarb
CO2 absorption efficiency (carbonator), %
Etot
CO2 capture rate (total), %
F
molar flow, mol/s
RCl
chlorine retention rate, %
T
temperature, °C
u0
superficial gas velocity, m/s
Ws
solid inventory, kg/m²
X
molar carbonate content, %
x
concentration, wt.%
y
concentration, vol.%
Abbreviations
ASU
air separation unit
BA
bottom ash
CaL
calcium looping
CCS
carbon capture and storage
CFB
circulating fluidized bed
FA
fly ash
FB
fluidized bed
FG
flue gas
FTIR
fourrier transformation infrared
GPU
gas processing unit
ICP-OES
inductively coupled plasma optical emission spectrometry
IGCC
integrated gasification combined cycle
IR
infrared
LHV
lower heating value
MSW
municipal solid waste
SRF
solid recovered fuel 2
TRL
technology readiness level
WtE
waste-to-energy
XRF
x-ray fluorescence
Subscripts
0
Ca-species in the limestone make-up
abs
absorbed
carb
carbonator
calc
calciner
eq
equilibrium
in
inlet
out
outlet
R
Ca-species circulating to the carbonator
SRF
solid recovered fuels
Vol
volatile matter
Greek letters
Φ
calciner CO2 concentration ratio, %
3
1
Introduction
The large-scale implementation of carbon capture and storage (CCS) processes in carbon intense industries is an efficient approach to limit the emissions of anthropogenic CO2 [1,2]. The calcium looping (CaL) process is a 2nd generation post-combustion CO2 capture process that is particularly suited for retrofitting existing power and industrial plants (e.g. cement plants) to reduce CO2 emissions [3,4]. The CaL process bases on the reversible carbonationcalcination reaction (Eq. 1) of limestone and CO2 [5]. 𝐶𝑎𝑂(𝑠) + 𝐶𝑂2(𝑔)
𝐶𝑎𝐶𝑂3(𝑠); ∆𝐻 = ± 178.2 𝑘𝐽/𝑚𝑜𝑙
(1)
The sorbent circulates between two interconnected fluidized bed (FB) reactors typically named carbonator and calciner. In the most mature CaL process configuration, carbonator and calciner are designed as circulating fluidized bed (CFB) reactors, that allows a compact reactor layout [6,7]. Figure 1 shows a schematic of the CaL process. CO2 (GPU)
CO2-depleted flue gas
Carbonator
Flue gas, FCO2 (host plant)
CaO + CaCO3
Calciner
CaO, FR
Purge
Fuel Limestone, F0
O2 (ASU)
Figure 1: Schematic of the CaL process. The main part of the CO2 present in the flue gas stream of the upstream unit is exothermically absorbed within the carbonator at approximately 650 °C. This temperature represents an ideal compromise between reaction kinetics and chemical equilibrium limitations of the carbonation reaction. The CO2-depleted flue gas is released to the environment, whereas the partly carbonated solid stream is transferred to the calciner. In the calciner, the temperature of the sorbent is raised close to 900 °C in order to achieve calcination conditions. A highly concentrated CO2 stream is available at the outlet of the calciner for further utilization or longterm storage after being purified in a gas process unit (GPU). The regenerated sorbent is looped back to the carbonator. The heat required in the calciner for the endothermic calcination reaction and for increasing the temperature of gas and solid phase, is supplied by means of oxy-fuel combustion of supplementary fuel [8]. The technically-pure oxygen that replaces the air-oxygen 4
is provided by on-site air separation unit (ASU). In order to account for the chemical deactivation and the accumulation of inert components (e.g. ash, CaSO4) in the circulating sorbent, a constant flux of fresh limestone is fed to process, while a purge stream is discharged from the solid loop. The CaL process characteristics are commonly discussed by the molar flows of the main material streams in the system (see Figure 1). Which are, the molar flow of CO2 that is fed to the carbonator with the flue gas of the host unit, FCO2, the molar flow of Ca-species that circulates between calciner and carbonator, FR, and the molar flow of fresh limestone in the make-up that is introduced to the system, F0. Due to the high operation temperature of carbonator (Tcarb ~ 650 °C) and calciner (Tcalc ~ 900 °C), an efficient heat utilization by means of a dedicated water-steam cycle is feasible. Thus, the CaL process offers relatively low efficiency penalties [9,10] and moderate CO2 avoidance costs [3,11] compared with other CO2 capture technologies, e.g. oxy-fuel combustion, integrated gasification combined cycle (IGCC) or amine-scrubbing [12,13]. Within the past years, the CaL process has undergone a rapid development, and numerous experimental investigations have been carried out. This advancement is sustained by the similarities between relevant components of the CaL process (e.g. CFB reactors, loop seals) and the already established state-of-the-art CFB combustion systems. Worldwide, several CaL test rigs are in operation that vary in terms of process layout, reactor design and equivalent thermal size. Small-scale units with an equivalent thermal size below 200 kWth are particularly suited for the investigations of CaL process fundamentals such as sorbent deactivation, sorbent attrition and hydrodynamic stability under laboratory conditions [14-17]. In order to further demonstrate the feasibility of the CaL process also under semi-industrial conditions, larger units with an equivalent thermal size over 200 kWth were built and operated [18-20]. Thus, the CaL process is currently ranked on a technology readiness level (TRL) of six due to the successful demonstration in relevant environment and scale [8]. As the next step in the course of the development of the CaL process, a demonstration plant in the size of 20 MWth is discussed [21,22]. The type of supplementary fuel that is utilized in the calciner has a significant influence on the CaL process performance due to the direct contact of circulating sorbent and fuel species (e.g. sulfur, ash) in the calciner [23]. Until now, fossil energy carriers such as hard coal, lignite or natural gas were mainly considered as supplementary fuel within the calciner [24,25]. A more sustainable attempt is the utilization of waste derived fuels. Thereby, net negative CO2 emissions are feasible due to the capture and long-term storage of biogenic based CO2 [26]. 5
Moreover, low or even negative fuel prices of waste derived fuels lead to significantly improved economics. Solid recovered fuels (SRF) represents a specially prepared type of fuel that is processed from production-specific wastes or municipal solid wastes [27,28]. Fuel-induced sorbent deactivation is caused by the accumulation of fuel-ash or calcium sulphate (CaSO4). Calcium sulphate is formed either directly (Eq. 2) or indirectly (Eq. 3) and remains stable under typical CaL process conditions (T < 1000 °C). 1 𝐶𝑎𝐶𝑂3(𝑠) + 𝑆𝑂2(𝑔) + 𝑂2(𝑔)→ 𝐶𝑎𝑆𝑂4(𝑠) + 𝐶𝑂2 2
(2)
1 𝐶𝑎𝑂(𝑠) + 𝑆𝑂2(𝑔) + 𝑂2(𝑔)→ 𝐶𝑎𝑠𝑂4(𝑠) 2
(3)
Indirect sulfation occurs, if the CO2 partial pressure in the reactor is less than the chemical equilibrium pressure of CO2 in the CaO-CO2-CaCO3 system (Eq. 1), whereas direct sulfation proceeds, if the CO2 partial pressure in the reactor is above the chemical equilibrium of CO2 [29]. Consequently, indirect sulfation dominates at the conditions typically prevailing in the calciner. When utilizing SRF in the CaL process, the smooth and continuous operation of the interconnected CFB system and the effects of SRF fuel species (e.g. Cl) on sorbent deactivation are of particular interest. Thus, the reaction mechanism of Ca-species with HCl are of additional concern. It is known that the absorption of HCl by limestone under typical conditions of FB boilers can be described by Eq. 4 and 5 [30,31]. 𝐶𝑎𝐶𝑂3(𝑠) + 2𝐻𝐶𝑙(𝑔)↔ 𝐶𝑎𝐶𝑙2(𝑠,𝑙) + 𝐻2𝑂(𝑔) + 𝐶𝑂2(𝑔)
(4)
𝐶𝑎𝑂(𝑠) + 2𝐻𝐶𝑙(𝑔)↔ 𝐶𝑎𝐶𝑙2(𝑠,𝑙) + 𝐻2𝑂
(5)
The HCl absorption phenomena is multi-layered, highly complex and mainly dependent on reaction temperature, prevailing gas atmosphere and the Cl/Ca molar ratio [32-34]. Several studies indicate the ability of Cl-retention by Ca-species during the combustion of waste derived fuels in FB combustion systems [35,36]. Under typical conditions of the oxy-fuel CaL calciner, the reaction product tends to form molten phases, which significantly alters the abovementioned reaction mechanism [37]. Until now, the HCl absorption behavior under typical condition of a CaL calciner has not been investigated yet. However, the influence of HCl absorption on the sorbent activity within the CaL process was assessed by means of thermogravimetric analysis. It was found, that the presence of HCl during the carbonation step 6
might lead to a more active sorbent due to the formation of larger pores during carbonation [38]. In another study, a fixed bed reactor system was applied to expose Ca-species to HCl during the cyclic CO2 absorption. Thereby, the discontinuously addition of HCl during the carbonation step was identified as a method to improve the CO2 capture capacity [39]. This work presents and discusses the performance of an SRF-fired CaL process due to extensive long-term pilot testing and evaluation for the first time. The experimental test runs were conducted at the 1 MWth CaL pilot plant at Technische Universität Darmstadt. Two different types of SRF were burnt in an oxy-fired CFB calciner, while CO2 was captured from a flue gas stream, which had a CO2 concertation similar to the flue gas of typical WtE plants. While another work focuses on the evaluation of the carbonator CO2 absorption performance [40], this study emphasis the calciner performance and the fate of chlorine within the CaL system. Moreover, a long-term data plot is shown for each type of SRF, respectively.
2 2.1
Experimental Section 1 MWth CaL Pilot Plant
The experimental investigations were conducted at the semi-industrial CaL pilot plant at Technische Universität Darmstadt. Figure 2 presents a simplified flowsheet of the 1 MWth CaL pilot plant. The carbonator riser has an inner diameter of 590 mm and a height of 8.6 m. The calciner riser has a height of 11.3 m and an inner diameter of 400 mm.
Solid stream
Solid sampling Gas analysis
Gaseous stream Calc II
Carb II
Calc III
Fly ash
Fly ash
CC
CaCO3
SRF
CFB Calciner
Combustion Chamber
CFB Carbonator
Coal
Air
CO2 Calc I
Air
Carb I
Air Bottom ash
Bottom ash
O2
Figure 2: Simplified flowsheet of the 1 MWth CaL pilot plant at Technische Universität Darmstadt. 7
In order to establish realistic experimental conditions, an on-site combustion chamber provides the flue gas to be decarbonized in the carbonator. Additionally, there is the possibility to introduce technically pure CO2 into the carbonator primary gas, which in fact allows investigations based on various CO2 concentrations or CO2 enriched ambient air (i.e. synthetic flue gas). The inlet temperature of the primary gas stream can be raised by means of electrical gas preheaters up to approximately 350 °C. The primary gas is then fed to the carbonator riser through the nozzle grid at the bottom of the reactor. The temperature of the carbonator can be adjusted by means of five axially arranged cooling tubes attached to the top of the reactor. The quantity of heat being removed, depends on the immerse depth of cooling tubes inside the carbonator riser. Carbonator bottom ash is discharged by means of a water-cooled screw conveyor attached to the solid material extraction pipe below the carbonator nozzle grid. The CO2-lean flue gas is released to the environment after heat removal and subsequent particle cleanup downstream of the carbonator cyclone. The partly carbonated particles are transported to the calciner either by means of a mechanically controlled screw conveyor or by means of a non-mechanically transport L-Valve. The solid stream is then introduced to the calciner return leg. At this point, the globally circulating solids are mixed together with the internally circulating solids from the calciner and the SRF. This mixture is introduced to the calciner riser, subsequently. The combustion gas of the calciner consist of technically pure oxygen (yO2 > 99 vol. %) and recirculated off-gas that is extracted downstream the calciner particle filter. The dilution of oxygen is required to avoid any hot spots within the fluidized bed. Similar to the carbonator, the calciner combustion gas can be preheated up to approximately 350 °C. The pre-heated combustion gas is fed to the calciner riser via the calciner nozzle grid and via the two opposite secondary air joints at approximately 2.9 m above the nozzle grid. The CO2-rich calciner off-gas undergoes heat removal and particle cleanup. Thereafter, part of the flue gas is recirculated back to the calciner inlet, the other part is released to the environment. The solid stream separated in the calciner cyclone is partly recirculated back to the calciner riser, whereas the other part is fed to the carbonator. The split ratio of the solid stream in the loop seal is controlled by means of a cone valve. Bottom ash is discharged from the calciner via a water-cooled extraction screw attached to the solid material extraction pipe at the calciner nozzle grid. The limestone make-up is fed to the return leg of the carbonator. The make-up dosing system consists of a supply vessel and a gravimetric dosing vessel attached underneath. The desired mass flow of limestone is adjusted by the rotating speed of the limestone feeding screw. The SRF feeding system consists of a supply bunker and a subsequently arranged gravimetric dosing 8
system. In order to avoid any leakage of process gases to the environment, two consecutively arranged rotary valves are installed at the process connection point at the calciner return leg. The space between the rotary valves is constantly flushed by inert gas in order to achieve a positive pressure gradient towards the process. Additionally, the smooth and stable SRF feed to the calciner riser is sustained by a constant flow of transport gas. Both CFB reactors are fully refractory lined, in order to minimize heat losses. Furthermore, the major components of the pilot plant are similar or equal to industrial standard, which means in particular, propane-fired start-up burner for each CFB reactor, two-pass heat exchangers and bag filters. The CFB reactor system is equipped with extractive gas analyzers at crucial process positions. Table 1 provides a summary of the applied gas measurement equipment at the 1 MWth CaL pilot plant. All gas extraction positions are also marked in Figure 2. Table 1: Online gas measurement equipment at the 1 MWth CaL pilot plant. Position CC, Carb II Carb I, Calc I Calc II Calc III IR, Paramagnetic, Infrared, IR, FTIR, Technique Psychometric Psychometric Paramagnetic Paramagnetic Species
O2, CO2, CO, SO2, H2O, NO
CO2, O2
HCl, O2, CO2, CO, NO, SO2
H2O
In addition to the continuous online gas measurements, the pilot plant is equipped with temperature, pressure and flow measurement at relevant process positions. Solid samples were taken during the experiment at all crucial points of the CaL process (see Figure 2). Relevant solid samples were analyzed in terms of X-ray fluorescence (XRF), inductively coupled plasma optical emission spectrometry (ICP-OES) and potentiometric titration to determine the chemical composition and the mass fractions of chlorine and heavy metals. 2.2
Materials
Two different types of SRF were used during the experimental long-term tests. In Table 2, the composition of the two types of SRF are given. The two SRF types were derived from German non-hazardous industrial waste, and mainly consist of plastics, paper, cardboard and textiles [41]. The fuels were fed to the process in form of industrial fluff. Throughout the investigations, the CaL system was supplied by a constant make-up flow of natural limestone. The limestone originates from a German limestone quarry. Table 3 shows the chemical composition of the natural limestone [42]. The particle diameter ranges from 25 to 500 µm with an average particle diameter of 180 µm. 9
Table 2: Composition of the solid recovered fuels. Parameter LHV Max. particle size Ash C H N S O Cl
Unit MJ/kg mm wt.%db wt.%db wt.%db wt.%db wt.%db wt.%db wt.%db
SRF I 21.4 30.0 11.0 58.0 8.20 0.54 0.22 21.3 0.73
SRF II 15.6 50.0 19.1 47.2 6.50 1.20 0.36 24.7 0.91
Table 3: Chemical composition of the limestone make-up. Species Mass fraction, wt.%
2.3
SiO2 0.50
Al2O3 0.10
Fe2O3 0.10
MgCO3 1.30
CaCO3 98.0
Evaluation parameter
Based on the data observed during the experiment, the following key evaluation parameters were determined. One important parameter for the CaL process is the CO2 absorption efficiency in the carbonator, Ecarb. It is derived from the amount of CO2 that is absorbed by the solid phase in the carbonator, FCO2,abs, according to Eq. 6.
𝐸𝑐𝑎𝑟𝑏 =
𝐹𝐶𝑂2,𝑎𝑏𝑠 𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑖𝑛
=
𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑖𝑛 ― 𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑜𝑢𝑡 𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑖𝑛
(6)
Thereby, FCO2,carb,in is the molar flow of CO2 that is fed to the carbonator, and FCO2,carb,out is the molar flow of CO2 that remains in the gas stream at the carbonator outlet. The maximum CO2 absorption efficiency is theoretically limited by the chemical equilibrium concentration of CO2, which mainly depends on the prevailing temperature in the carbonator [43]. In addition to the CO2 absorption efficiency of the carbonator, the total CO2 capture efficiency of the system, Etot, is calculated according to Eq. 7.
𝐸𝑡𝑜𝑡 =
𝐹𝐶𝑂2,𝑐𝑎𝑙𝑐,𝑜𝑢𝑡 𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑖𝑛 + 𝐹0 + 𝐹𝐶𝑂2,𝑆𝑅𝐹
(7)
Hereby, F0 is the molar flux of CO2 that is released by the first calcination of the fresh limestone make-up, FCO2,SRF is the CO2 that is released during the oxy-fuel combustion of SRF in the calciner. Due to the fact, that the CO2, which is formed in the oxy-fuel calciner is captured 10
nearly completely, the total CO2 capture efficiency is always higher than the CO2 absorption efficiency of the carbonator. Similar to the CO2 absorption efficiency of the carbonator, the calciner efficiency, Ecalc, is a crucial metric in the CaL process. It is defined according to Eq. 8.
𝐸𝑐𝑎𝑙𝑐 = 1 ―
𝑋𝑐𝑎𝑙𝑐
(8)
𝑋𝑐𝑎𝑟𝑏
Hereby, Xcalc is the molar carbonate content at the outlet of the calciner, and Xcarb is the molar carbonate content at the inlet of the calciner.
3
Results and Discussion
3.1
Overview
The overall goal was to demonstrate the feasibility of continuous and stable CO2 capture while using SRF as supplementary heat source. Throughout two consecutive experimental test campaigns, the pilot plant was operation more than 230 hours in coupled CFB mode under realistic CaL process conditions. This particularly implies the oxy-fuel combustion of SRF in the CFB calciner. Table 4 summarizes the range of conditions prevailing during the experimental tests. Table 4: Range of experimental test conditions. System
Carbonator
Calciner
Global
Parameter Solid inventory Superficial gas velocity Temperature CO2 conc. in primary gas CO2 absorption efficiency Solid inventory Superficial gas velocity Temperature O2 conc. in primary gas Specific sorbent circulation Specific molar make-up feed Total CO2 capture rate
Symbol Ws,carb u0,carb Tcarb yCO2,carb,in Ecarb Ws,calc u0,calc Tcalc yO2,calc,in FR/FCO2 F0/FCO2 Etot
Unit kg/m² m/s °C vol.% % kg/m² m/s °C vol.%dry mole% mole% %
Value 300 - 700 2.5 - 3.5 610 - 690 9.0 - 10.5 60 - 85 150 - 400 4.5 - 5.5 750 - 900 40 - 65 7.5 - 15 0.09 - 0.35 82 - 94
CO2 absorption efficiencies close to chemical equilibrium conditions were reached in the carbonator, while the total CO2 capture efficiency varies between 82 and 94 %. Thus, for the
11
first time, the CaL process was successfully fueled by oxy-fuel combustion of a waste derived fuel. Figure 3 shows an exemplary data plot of relevant CaL process parameters of the carbonator (Graph a) and the calciner (Graph b) for a period of six hours. During this period, the CaL process was oxy-fueled by SRF I. 100
800
a)
70
720
60 680 50
Ecarb
40 100
yi (vol.%)dry
Tc arb (°C)
760 80
Etot
Tcarb 640 900
b)
80
880
60
860
40
840
20
820 yO2,in
0
Tcalc (°C)
Ecarb , Etot (%)
90
0
1
2
3
yCO2,out 4
Tcalc 5
6
800
t (h) Figure 3: Exemplary data plot of key process parameter for carbonator (Graph a), and calciner (Graph b) while firing SRF I (F0/FCO2 ~ 0.12, FR/FCO2 ~ 11.1). Approximately 75 % of the CO2 contained in the flue gas was absorbed in the carbonator, which results in a total CO2 capture efficiency above 90 %. The solid inventory induced pressure drop was between 40 and 60 mbar in the carbonator and between 15 and 25 mbar in the calciner respectively. This corresponds to an average specific solid inventory of approximately 580 kg/m² and 280 kg/m² for carbonator and calciner, respectively. The calciner was operated within a temperature range of 820 - 850 °C, and with an inlet oxygen concentration slightly above 60 vol.%dry. During this period, the oxygen excess at the calciner outlet was between 4 and 13 vol%dry (average: 7.3 vol.%dry). This large oxygen excess was required to ensure a complete conversion of the volatile matter of SRF (xVol.,SRFI: 70 - 80 wt.%ar) at all time, while the fuel was inhomogeneous and consequently, fuel mass flow was changing considerably. In 12
addition to the high oxygen excess, the CO2 product stream at outlet of the calciner was further diluted by nitrogen required for the loop seal fluidization and the flushing of the fuel feeding system. The temperature fluctuations of the calciner were mainly caused by the inhomogeneous nature of the SRF. The carbonator was operated slightly below 680 °C. Under these circumstances, no active reactor cooling was required. This fact was caused by the lower CO2 concentration in the flue gas to be decarbonized in the carbonator. In contrast to coal-fired power plants, the CO2 concentration in the flue gas of a WtE plant is reduced by almost 30 %. Thus, less heat was released by the exothermic carbonation reaction in the carbonator. Figure 4 shows a similar data plot presenting the same data set for the carbonator (Graph a) and for the calciner (Graph b) while oxy-firing SRF II. 100
800
a)
760 80 70
720
60
Tc arb (°C)
Ecarb , Etot (%)
90
680 50
Ecarb
Etot
Tcarb
40 100
640 900
80
870
60
840
40
810
20
780 yO2,in
0
Tcalc (°C)
yi (vol.%)dry
b)
0
0.5
1
1.5
2
yCO2,out 2.5
Tcalc 3
750 3.5
t (h) Figure 4: Exemplary data plot of key process parameter for the carbonator (Graph a) and calciner (Graph b) while firing SRF II (F0/FCO2 ~ 0.22, FR/FCO2 ~ 12.2). The values for CO2 absorption efficiency and the total CO2 capture efficiency achieved during that test period are almost similar to the aforementioned results (Ecarb ~ 75 %, Etot ~ 90 %). In contrast, the overall plant operability was noteworthy different during this period. The coarse size specification of SRF II (< 50 mm) led to the accumulation of relatively large ash fractions, such as metal, glass or mineral matter in the solid looping system. This, in fact negatively 13
influences the hydrodynamics of the interconnected CFB reactors. Exemplarily, the rapid drop in the CO2 absorption efficiency within the carbonator at 0.75 h was caused by defluidization in the solid looping system due to the accumulation of coarse ash in the calciner loop seal. In order to prevent the accumulation of coarse ashes, the make-up flux was increased by almost 80 % in contrast to the operational period firing SRF I. Consequently, extraction of bottom ash was carried out more frequently. The need for a relatively large oxygen excess at the calciner outlet was also due for SRF II (xVol.,SRFII: 60 - 75 wt.%ar). 3.2
Calciner performance
The calciner efficiency (Eq. 8) is mainly depended on the calciner active space time, the prevailing CO2 concentration and the calciner temperature. The last two parameters are directly linked by the chemical equilibrium of the CaO-CO2-CaCO3 system. Accordingly, calciner temperature close to 900 °C are required to achieve a complete sorbent calcination under ideal conditions of an oxy-fired CaL calciner. At the 1 MWth CaL pilot plant, dilution of the CO2 product stream in the calciner occurred due to practical reasons such as flushing of the various pressure measurements, fluidization of the calciner loop seal and the sealing gas required in the fuel feeding system. Moreover, the oxygen excess at the calciner outlet was higher compared to previous experimental investigations of the CaL process based on hard coal or lignite as supplementary fuel at the same pilot plant [19,23,44] (see Chapter 3.1). Therefore, a sufficient sorbent calcination was feasible even at a calciner operation temperature in the range of 850 880 °C. In order to describe the calciner characteristic with regard to the chemical equilibrium conditions, the ratio, Φ, of the CO2 concentration at the outlet, yCO2,calc,out, and at the chemical equilibrium, yCO2,eq, is introduced:
𝛷=
𝑦𝐶𝑂2,𝑐𝑎𝑙𝑐,𝑜𝑢𝑡
(9)
𝑦𝐶𝑂2,𝑒𝑞
The CO2 concentration at chemical equilibrium is calculated by Eq. 10 [45]:
𝑦𝐶𝑂2,𝑒𝑞 = 4.137 ∗ 107𝑒
―
20474 𝑇𝑐𝑎𝑙𝑐
(10)
For a sufficient sorbent calcination, Φ needs to be < 1, whereas a value > 1 might lead to slow calcination or re-carbonation. These parameters are shown in Figure 5 for representative experimental conditions depended on the average calciner temperature. 14
100
2
1.5
60 1 40 0.5
20 Ecalc 0 800
Φ (-)
Ecalc (%)
80
820
840
860
880
Φ 0 900
Tcalc (°C) Figure 5: Calciner efficiency (Ecalc) and CO2 concentration ratio (Φ) depended on calciner temperature (Tcalc). Under the experimental conditions, the ratio of CO2 concentration during operation was between 0.53 and 1.75, respectively. The calciner efficiency was between 9.6 and 97 %. It can be seen that an average calciner temperature over 850 °C lead to CO2 concentration ratios below one, which results in a nearly complete sorbent calcination (Ecalc > 95 %). Once the temperature of the calciner was below 830 °C, sorbent calcination was insufficient (Ecalc < 40 %).
3.3
The Fate of Chlorine
Chlorine is introduced to the CaL process via the various waste fractions contained in the SRF. Understanding the fate of chlorine within the CaL system is crucial for an appropriate design of subsequent heat recovery systems and gas cleaning units. In this regard, any form of chlorineinduced corrosion of boiler materials is of major concern [46-48]. During oxy-fuel combustion in the CFB calciner, chlorine is released to the gas phase via HCl and partly absorbed by the solid phase. Part of solid phase bound chlorine is transported to the carbonator by the circulating sorbent. Thus, chlorine is released by the CaL process via gaseous and solid streams from the carbonator and from the calciner, respectively. The ranges of chlorine mass fraction and the total mass flow of the solid effluent streams are summarized in Table 5.
15
Table 5: Range of chlorine mass fraction and mass flow of relevant solid effluent streams from the CaL process. (BA: bottom ash, FA: fly ash). Parameter Mass flow, kg/h Mass fraction, wt.%
BACarb 2.5 - 38 0.20 - 0.32
BACalc 5.0 - 58 0.21 - 0.36
FACarb 1.0 - 21 1.61
FACalc 5.8 - 44 2.2 - 3.3
The chemical solid analysis shows that chlorine concentration is highest in the calciner fly ash, while carbonator fly ash shows a lower value in the same order of magnitude. In contrast to that, the chlorine mass fraction in bottom ashes discharged from carbonator and calciner are significantly lower. This phenomena is mainly caused by the fine Ca-particles present in the fly ash (dp,50 < 50 µm), that favor the absorption of gaseous HCl. Moreover, fly ash particles are additionally exposed to the calciner flue gases during the course of heat removal and in the bag filter. Thus, further HCl absorption is likely to occur in this process sections. Based on the chlorine introduced by SRF and the solid and gaseous effluent streams from the CaL process, a chlorine balance was established for a representative operation point (Table 6). Table 6: Chlorine balance for a representative CaL operation point2 (Tcarb: 679 °C, Tcalc: 841 °C, Etot: 83.2 %, Pth,calc: 550 kWth, FCl,SRFI: 5.42 mol/s). Parameter Cl molar flow, mol/s Cl share, %3 Cl share, %4
BACarb 0.283 0.412 0.510
BACalc 1.80 2.62 3.24
FACarb 1.61 2.34 2.89
FACalc 57.9 84.2 104
FGCalc 7.01 10.5 12.9
Sum 68.6 100 123
The fly ash of the calciner represents the major chlorine effluent from the CaL process. Moreover, it can be seen, that there is a deviation of 23 % between the chlorine feed in the SRF and the chlorine present in the effluent streams. This might be caused by several reasons, such as the fluctuation of the actual chlorine present in the SRF or the fact that the extracted solid samples represent only a minor share compared to the corresponding mass flows. Furthermore, it is likely, that HCl absorption from the calciner flue gas continues in the course of the flue gas path downstream the calciner cyclone. Due to the fact, that HCl was measured directly at the outlet of the calciner cyclone, chlorine effluent in the calciner flue gas tends to be overestimated. Knowing the concentration of HCl in the calciner flue gas, the chlorine retention rate, RCl, is calculated according to Eq. 11. This method assumes, that all gaseous chlorine in the calciner
1
Only one representative solid sample was available for the carbonator fly ash.
2
HCl in the carbonator flue gas was neglected.
3
This share refers to the sum of all chlorine in the considered effluent
4
This share refers to the chlorine in the SRF 16
off-gas is present as HCl, moreover, solid phase bound chlorine remains inert within the carbonator. Thus, there is no gaseous Cl effluent in the carbonator off-gas.
𝑅𝐶𝑙 = 1 ―
𝐹𝐻𝐶𝑙,𝑐𝑎𝑙𝑐,𝑜𝑢𝑡
(11)
𝐹𝐶𝑙,𝑆𝑅𝐹
Hereby, FHCl,FG is the molar flow of HCl in the calciner flue gas downstream of the extraction point of the recirculation gas and, FCl,fuel is the molar flow of chlorine introduced by the SRF. Figure 6 shows the average chlorine retention rate for stable operation points of one hour, each. It can be seen that the chlorine retention rate was over 82 % during all experimental test points, respectively. Furthermore, there is a tendency of an increasing chlorine retention rate along with increasing calciner temperatures. While comparing the two types of SRF, there is no distinguished difference in terms of Cl retention. As already mentioned in the course of the chlorine balance, HCl was measured directly at the outlet of the calciner cyclone. Thus, the likely further HCl absorption by Ca-rich calciner fly ash in the course of the calciner flue gas treatment path was not considered here. Therefore, chlorine retention rates tends to be even higher than those shown in Figure 6. 100 95
RCl (%)
90 85 80 SRF I SRF II
75 70 800
820
840
860
880
900
Tcalc (°C) Figure 6: Chlorine retention rate (RCl) dependent on the average calciner temperature (Tcalc).
4
Conclusions
For the first time, the CaL process has been operated while a waste derived fuel was oxy-fired in the calciner. Two types of SRF were fed to a 1 MWth pilot plant in the form of raw fluff. During more than 230 h of coupled CFB operation, CO2 was continuously captured from a flue 17
gas steam characterized by a CO2 concentration similar to typical WtE plants fueled by MSW. Under these conditions, carbonator CO2 absorption rates in the range of 80 - 85 % and total CO2 capture efficiencies over 90 % were achieved at representative CaL process conditions. It was found, that fuel quality has a significant influence on the operability of the interconnected CFB system. The pilot plant operability when using SRF I was relatively smooth and stable. When using SRF II, the enrichment of coarse ash fractions within the solid phase was the main reason that negatively affect the hydrodynamics of the interconnected CFB system. For a stable operation point, a chlorine balance reveals, that more than 80 % of the fuel-chlorine is released in the fly ash of the calciner. Due to the limestone based reactor inventory, the retention rate of chlorine was over 82 % throughout the experimental test points. It was found, that a calciner operation temperature above 860 °C is desirable in order to maximize Cl retention. When taking into account the organic waste fractions typically contained in SRF and MSW, this work successfully demonstrates the feasibility of net negative CO2 emissions by means of the CaL process at semi-industrial scale.
Acknowledgements The research leading to these results has received funding from the German Ministry of Economic Affairs and Energy based on a resolution of the German Parliament (MONIKA, FKZ: 03ET7089) and the SUEZ Group. The funding is gratefully acknowledged.
18
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Martin Haaf: Methodology, Investigation, Writing - Original Draft, Writing - Review & Editing Jens Peters: Investigation, Methodology Jochen Hilz: Investigation Antonio Unger: Resources, Supervision Jochen Ströhle: Project administration, Writing - Review & Editing, Supervision Bernd Epple: Supervision, Funding acquisition
Declaration of interests ☒ 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. ☐The authors declare the following financial interests/personal relationships which may be considered as potential competing interests:
Oxy-fuel combustion of solid recovered fuels in the calcium looping (CaL) process. Over 230 h of representative CaL operation at a 1 MWth pilot plant. Total CO2 capture rates > 90 % have been achieved. Assessment of the fate of chlorine within the calcium looping system.
24
S0894-1777(19)31925-9 https://doi.org/10.1016/j.expthermflusci.2019.110023 ETF 110023
To appear in:
Experimental Thermal and Fluid Science
Received Date: Revised Date: Accepted Date:
12 November 2019 11 December 2019 11 December 2019
Please cite this article as: M. Haaf, J. Peters, J. Hilz, A. Unger, J. Ströhle, B. Epple, Combustion of Solid Recovered Fuels within the Calcium Looping Process - Experimental Demonstration at 1 MWth Scale, Experimental Thermal and Fluid Science (2019), doi: https://doi.org/10.1016/j.expthermflusci.2019.110023
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Combustion of Solid Recovered Fuels within the Calcium Looping Process - Experimental Demonstration at 1 MWth Scale Martin Haaf1*, Jens Peters1, Jochen Hilz1, Antonio Unger2, Jochen Ströhle1, Bernd Epple1 1Institute
for Energy Systems and Technology, Technische Universität Darmstadt, Otto-BerndtStraße 2, 64287 Darmstadt, Germany
2SUEZ
Deutschland GmbH, Binnenhafenstraße 19, 68159 Mannheim, Germany
*Corresponding Author: Martin Haaf, Institute for Energy Systems and Technology, Technische Universität Darmstadt Otto-Berndt-Straße 2, 64287 Darmstadt, Germany Tel.: +49 6151 / 16 22675 Fax: +49 6151 / 1622690 Email: [email protected] Abstract
The calcium looping (CaL) process is an efficient post-combustion CO2 capture technology based on the reversible carbonation-calcination reaction of natural lime. This work presents the results obtained during long-term operation at the 1 MWth CaL pilot plant at Technische Universität Darmstadt. During more than 230 h of representative CaL operation, the calciner was heated by oxy-fuel combustion of waste derived fuels. During the experimental investigation, two types of solid recovered fuels (SRF) were utilized. Both types of SRF were fed to the process in the form of raw fluff, similar to typical industrial applications. The flue gas to be decarbonized in the carbonator, was supplied by an on-site combustion chamber. The CO2 concentration was kept between 9.5 and 10.5 vol.% as typical value for waste-to-energy (WtE) plants fueled by municipal solid waste (MSW). Over a wide range of operation conditions, CO2 absorption rates of 80 % and total CO2 capture rates over 90 % were achieved. Within this work, exemplary long-term data plots of relevant CaL process parameters are shown for each type of SRF. A chlorine balance for a representative operation period is established based on the relevant effluent streams from the CaL system. Thereby, it was found that the calciner fly ash represents the major chlorine effluent. Furthermore, the retention rate of chlorine was above 82 % throughout all test points, respectively. When taking into account the organic waste fractions typically contained in SRF and MSW, this work successfully demonstrates the feasibility of net negative CO2 emissions by means of the CaL process at semiindustrial scale.
1
Keywords: calcium looping (CaL) process, carbon capture and storage (CCS), pilot plant, waste derived fuel, solid recovered fuels (SRF), negative emission technology (NET)
Symbols Ecalc
calciner efficiency, %
Ecarb
CO2 absorption efficiency (carbonator), %
Etot
CO2 capture rate (total), %
F
molar flow, mol/s
RCl
chlorine retention rate, %
T
temperature, °C
u0
superficial gas velocity, m/s
Ws
solid inventory, kg/m²
X
molar carbonate content, %
x
concentration, wt.%
y
concentration, vol.%
Abbreviations
ASU
air separation unit
BA
bottom ash
CaL
calcium looping
CCS
carbon capture and storage
CFB
circulating fluidized bed
FA
fly ash
FB
fluidized bed
FG
flue gas
FTIR
fourrier transformation infrared
GPU
gas processing unit
ICP-OES
inductively coupled plasma optical emission spectrometry
IGCC
integrated gasification combined cycle
IR
infrared
LHV
lower heating value
MSW
municipal solid waste
SRF
solid recovered fuel 2
TRL
technology readiness level
WtE
waste-to-energy
XRF
x-ray fluorescence
Subscripts
0
Ca-species in the limestone make-up
abs
absorbed
carb
carbonator
calc
calciner
eq
equilibrium
in
inlet
out
outlet
R
Ca-species circulating to the carbonator
SRF
solid recovered fuels
Vol
volatile matter
Greek letters
Φ
calciner CO2 concentration ratio, %
3
1
Introduction
The large-scale implementation of carbon capture and storage (CCS) processes in carbon intense industries is an efficient approach to limit the emissions of anthropogenic CO2 [1,2]. The calcium looping (CaL) process is a 2nd generation post-combustion CO2 capture process that is particularly suited for retrofitting existing power and industrial plants (e.g. cement plants) to reduce CO2 emissions [3,4]. The CaL process bases on the reversible carbonationcalcination reaction (Eq. 1) of limestone and CO2 [5]. 𝐶𝑎𝑂(𝑠) + 𝐶𝑂2(𝑔)
𝐶𝑎𝐶𝑂3(𝑠); ∆𝐻 = ± 178.2 𝑘𝐽/𝑚𝑜𝑙
(1)
The sorbent circulates between two interconnected fluidized bed (FB) reactors typically named carbonator and calciner. In the most mature CaL process configuration, carbonator and calciner are designed as circulating fluidized bed (CFB) reactors, that allows a compact reactor layout [6,7]. Figure 1 shows a schematic of the CaL process. CO2 (GPU)
CO2-depleted flue gas
Carbonator
Flue gas, FCO2 (host plant)
CaO + CaCO3
Calciner
CaO, FR
Purge
Fuel Limestone, F0
O2 (ASU)
Figure 1: Schematic of the CaL process. The main part of the CO2 present in the flue gas stream of the upstream unit is exothermically absorbed within the carbonator at approximately 650 °C. This temperature represents an ideal compromise between reaction kinetics and chemical equilibrium limitations of the carbonation reaction. The CO2-depleted flue gas is released to the environment, whereas the partly carbonated solid stream is transferred to the calciner. In the calciner, the temperature of the sorbent is raised close to 900 °C in order to achieve calcination conditions. A highly concentrated CO2 stream is available at the outlet of the calciner for further utilization or longterm storage after being purified in a gas process unit (GPU). The regenerated sorbent is looped back to the carbonator. The heat required in the calciner for the endothermic calcination reaction and for increasing the temperature of gas and solid phase, is supplied by means of oxy-fuel combustion of supplementary fuel [8]. The technically-pure oxygen that replaces the air-oxygen 4
is provided by on-site air separation unit (ASU). In order to account for the chemical deactivation and the accumulation of inert components (e.g. ash, CaSO4) in the circulating sorbent, a constant flux of fresh limestone is fed to process, while a purge stream is discharged from the solid loop. The CaL process characteristics are commonly discussed by the molar flows of the main material streams in the system (see Figure 1). Which are, the molar flow of CO2 that is fed to the carbonator with the flue gas of the host unit, FCO2, the molar flow of Ca-species that circulates between calciner and carbonator, FR, and the molar flow of fresh limestone in the make-up that is introduced to the system, F0. Due to the high operation temperature of carbonator (Tcarb ~ 650 °C) and calciner (Tcalc ~ 900 °C), an efficient heat utilization by means of a dedicated water-steam cycle is feasible. Thus, the CaL process offers relatively low efficiency penalties [9,10] and moderate CO2 avoidance costs [3,11] compared with other CO2 capture technologies, e.g. oxy-fuel combustion, integrated gasification combined cycle (IGCC) or amine-scrubbing [12,13]. Within the past years, the CaL process has undergone a rapid development, and numerous experimental investigations have been carried out. This advancement is sustained by the similarities between relevant components of the CaL process (e.g. CFB reactors, loop seals) and the already established state-of-the-art CFB combustion systems. Worldwide, several CaL test rigs are in operation that vary in terms of process layout, reactor design and equivalent thermal size. Small-scale units with an equivalent thermal size below 200 kWth are particularly suited for the investigations of CaL process fundamentals such as sorbent deactivation, sorbent attrition and hydrodynamic stability under laboratory conditions [14-17]. In order to further demonstrate the feasibility of the CaL process also under semi-industrial conditions, larger units with an equivalent thermal size over 200 kWth were built and operated [18-20]. Thus, the CaL process is currently ranked on a technology readiness level (TRL) of six due to the successful demonstration in relevant environment and scale [8]. As the next step in the course of the development of the CaL process, a demonstration plant in the size of 20 MWth is discussed [21,22]. The type of supplementary fuel that is utilized in the calciner has a significant influence on the CaL process performance due to the direct contact of circulating sorbent and fuel species (e.g. sulfur, ash) in the calciner [23]. Until now, fossil energy carriers such as hard coal, lignite or natural gas were mainly considered as supplementary fuel within the calciner [24,25]. A more sustainable attempt is the utilization of waste derived fuels. Thereby, net negative CO2 emissions are feasible due to the capture and long-term storage of biogenic based CO2 [26]. 5
Moreover, low or even negative fuel prices of waste derived fuels lead to significantly improved economics. Solid recovered fuels (SRF) represents a specially prepared type of fuel that is processed from production-specific wastes or municipal solid wastes [27,28]. Fuel-induced sorbent deactivation is caused by the accumulation of fuel-ash or calcium sulphate (CaSO4). Calcium sulphate is formed either directly (Eq. 2) or indirectly (Eq. 3) and remains stable under typical CaL process conditions (T < 1000 °C). 1 𝐶𝑎𝐶𝑂3(𝑠) + 𝑆𝑂2(𝑔) + 𝑂2(𝑔)→ 𝐶𝑎𝑆𝑂4(𝑠) + 𝐶𝑂2 2
(2)
1 𝐶𝑎𝑂(𝑠) + 𝑆𝑂2(𝑔) + 𝑂2(𝑔)→ 𝐶𝑎𝑠𝑂4(𝑠) 2
(3)
Indirect sulfation occurs, if the CO2 partial pressure in the reactor is less than the chemical equilibrium pressure of CO2 in the CaO-CO2-CaCO3 system (Eq. 1), whereas direct sulfation proceeds, if the CO2 partial pressure in the reactor is above the chemical equilibrium of CO2 [29]. Consequently, indirect sulfation dominates at the conditions typically prevailing in the calciner. When utilizing SRF in the CaL process, the smooth and continuous operation of the interconnected CFB system and the effects of SRF fuel species (e.g. Cl) on sorbent deactivation are of particular interest. Thus, the reaction mechanism of Ca-species with HCl are of additional concern. It is known that the absorption of HCl by limestone under typical conditions of FB boilers can be described by Eq. 4 and 5 [30,31]. 𝐶𝑎𝐶𝑂3(𝑠) + 2𝐻𝐶𝑙(𝑔)↔ 𝐶𝑎𝐶𝑙2(𝑠,𝑙) + 𝐻2𝑂(𝑔) + 𝐶𝑂2(𝑔)
(4)
𝐶𝑎𝑂(𝑠) + 2𝐻𝐶𝑙(𝑔)↔ 𝐶𝑎𝐶𝑙2(𝑠,𝑙) + 𝐻2𝑂
(5)
The HCl absorption phenomena is multi-layered, highly complex and mainly dependent on reaction temperature, prevailing gas atmosphere and the Cl/Ca molar ratio [32-34]. Several studies indicate the ability of Cl-retention by Ca-species during the combustion of waste derived fuels in FB combustion systems [35,36]. Under typical conditions of the oxy-fuel CaL calciner, the reaction product tends to form molten phases, which significantly alters the abovementioned reaction mechanism [37]. Until now, the HCl absorption behavior under typical condition of a CaL calciner has not been investigated yet. However, the influence of HCl absorption on the sorbent activity within the CaL process was assessed by means of thermogravimetric analysis. It was found, that the presence of HCl during the carbonation step 6
might lead to a more active sorbent due to the formation of larger pores during carbonation [38]. In another study, a fixed bed reactor system was applied to expose Ca-species to HCl during the cyclic CO2 absorption. Thereby, the discontinuously addition of HCl during the carbonation step was identified as a method to improve the CO2 capture capacity [39]. This work presents and discusses the performance of an SRF-fired CaL process due to extensive long-term pilot testing and evaluation for the first time. The experimental test runs were conducted at the 1 MWth CaL pilot plant at Technische Universität Darmstadt. Two different types of SRF were burnt in an oxy-fired CFB calciner, while CO2 was captured from a flue gas stream, which had a CO2 concertation similar to the flue gas of typical WtE plants. While another work focuses on the evaluation of the carbonator CO2 absorption performance [40], this study emphasis the calciner performance and the fate of chlorine within the CaL system. Moreover, a long-term data plot is shown for each type of SRF, respectively.
2 2.1
Experimental Section 1 MWth CaL Pilot Plant
The experimental investigations were conducted at the semi-industrial CaL pilot plant at Technische Universität Darmstadt. Figure 2 presents a simplified flowsheet of the 1 MWth CaL pilot plant. The carbonator riser has an inner diameter of 590 mm and a height of 8.6 m. The calciner riser has a height of 11.3 m and an inner diameter of 400 mm.
Solid stream
Solid sampling Gas analysis
Gaseous stream Calc II
Carb II
Calc III
Fly ash
Fly ash
CC
CaCO3
SRF
CFB Calciner
Combustion Chamber
CFB Carbonator
Coal
Air
CO2 Calc I
Air
Carb I
Air Bottom ash
Bottom ash
O2
Figure 2: Simplified flowsheet of the 1 MWth CaL pilot plant at Technische Universität Darmstadt. 7
In order to establish realistic experimental conditions, an on-site combustion chamber provides the flue gas to be decarbonized in the carbonator. Additionally, there is the possibility to introduce technically pure CO2 into the carbonator primary gas, which in fact allows investigations based on various CO2 concentrations or CO2 enriched ambient air (i.e. synthetic flue gas). The inlet temperature of the primary gas stream can be raised by means of electrical gas preheaters up to approximately 350 °C. The primary gas is then fed to the carbonator riser through the nozzle grid at the bottom of the reactor. The temperature of the carbonator can be adjusted by means of five axially arranged cooling tubes attached to the top of the reactor. The quantity of heat being removed, depends on the immerse depth of cooling tubes inside the carbonator riser. Carbonator bottom ash is discharged by means of a water-cooled screw conveyor attached to the solid material extraction pipe below the carbonator nozzle grid. The CO2-lean flue gas is released to the environment after heat removal and subsequent particle cleanup downstream of the carbonator cyclone. The partly carbonated particles are transported to the calciner either by means of a mechanically controlled screw conveyor or by means of a non-mechanically transport L-Valve. The solid stream is then introduced to the calciner return leg. At this point, the globally circulating solids are mixed together with the internally circulating solids from the calciner and the SRF. This mixture is introduced to the calciner riser, subsequently. The combustion gas of the calciner consist of technically pure oxygen (yO2 > 99 vol. %) and recirculated off-gas that is extracted downstream the calciner particle filter. The dilution of oxygen is required to avoid any hot spots within the fluidized bed. Similar to the carbonator, the calciner combustion gas can be preheated up to approximately 350 °C. The pre-heated combustion gas is fed to the calciner riser via the calciner nozzle grid and via the two opposite secondary air joints at approximately 2.9 m above the nozzle grid. The CO2-rich calciner off-gas undergoes heat removal and particle cleanup. Thereafter, part of the flue gas is recirculated back to the calciner inlet, the other part is released to the environment. The solid stream separated in the calciner cyclone is partly recirculated back to the calciner riser, whereas the other part is fed to the carbonator. The split ratio of the solid stream in the loop seal is controlled by means of a cone valve. Bottom ash is discharged from the calciner via a water-cooled extraction screw attached to the solid material extraction pipe at the calciner nozzle grid. The limestone make-up is fed to the return leg of the carbonator. The make-up dosing system consists of a supply vessel and a gravimetric dosing vessel attached underneath. The desired mass flow of limestone is adjusted by the rotating speed of the limestone feeding screw. The SRF feeding system consists of a supply bunker and a subsequently arranged gravimetric dosing 8
system. In order to avoid any leakage of process gases to the environment, two consecutively arranged rotary valves are installed at the process connection point at the calciner return leg. The space between the rotary valves is constantly flushed by inert gas in order to achieve a positive pressure gradient towards the process. Additionally, the smooth and stable SRF feed to the calciner riser is sustained by a constant flow of transport gas. Both CFB reactors are fully refractory lined, in order to minimize heat losses. Furthermore, the major components of the pilot plant are similar or equal to industrial standard, which means in particular, propane-fired start-up burner for each CFB reactor, two-pass heat exchangers and bag filters. The CFB reactor system is equipped with extractive gas analyzers at crucial process positions. Table 1 provides a summary of the applied gas measurement equipment at the 1 MWth CaL pilot plant. All gas extraction positions are also marked in Figure 2. Table 1: Online gas measurement equipment at the 1 MWth CaL pilot plant. Position CC, Carb II Carb I, Calc I Calc II Calc III IR, Paramagnetic, Infrared, IR, FTIR, Technique Psychometric Psychometric Paramagnetic Paramagnetic Species
O2, CO2, CO, SO2, H2O, NO
CO2, O2
HCl, O2, CO2, CO, NO, SO2
H2O
In addition to the continuous online gas measurements, the pilot plant is equipped with temperature, pressure and flow measurement at relevant process positions. Solid samples were taken during the experiment at all crucial points of the CaL process (see Figure 2). Relevant solid samples were analyzed in terms of X-ray fluorescence (XRF), inductively coupled plasma optical emission spectrometry (ICP-OES) and potentiometric titration to determine the chemical composition and the mass fractions of chlorine and heavy metals. 2.2
Materials
Two different types of SRF were used during the experimental long-term tests. In Table 2, the composition of the two types of SRF are given. The two SRF types were derived from German non-hazardous industrial waste, and mainly consist of plastics, paper, cardboard and textiles [41]. The fuels were fed to the process in form of industrial fluff. Throughout the investigations, the CaL system was supplied by a constant make-up flow of natural limestone. The limestone originates from a German limestone quarry. Table 3 shows the chemical composition of the natural limestone [42]. The particle diameter ranges from 25 to 500 µm with an average particle diameter of 180 µm. 9
Table 2: Composition of the solid recovered fuels. Parameter LHV Max. particle size Ash C H N S O Cl
Unit MJ/kg mm wt.%db wt.%db wt.%db wt.%db wt.%db wt.%db wt.%db
SRF I 21.4 30.0 11.0 58.0 8.20 0.54 0.22 21.3 0.73
SRF II 15.6 50.0 19.1 47.2 6.50 1.20 0.36 24.7 0.91
Table 3: Chemical composition of the limestone make-up. Species Mass fraction, wt.%
2.3
SiO2 0.50
Al2O3 0.10
Fe2O3 0.10
MgCO3 1.30
CaCO3 98.0
Evaluation parameter
Based on the data observed during the experiment, the following key evaluation parameters were determined. One important parameter for the CaL process is the CO2 absorption efficiency in the carbonator, Ecarb. It is derived from the amount of CO2 that is absorbed by the solid phase in the carbonator, FCO2,abs, according to Eq. 6.
𝐸𝑐𝑎𝑟𝑏 =
𝐹𝐶𝑂2,𝑎𝑏𝑠 𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑖𝑛
=
𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑖𝑛 ― 𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑜𝑢𝑡 𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑖𝑛
(6)
Thereby, FCO2,carb,in is the molar flow of CO2 that is fed to the carbonator, and FCO2,carb,out is the molar flow of CO2 that remains in the gas stream at the carbonator outlet. The maximum CO2 absorption efficiency is theoretically limited by the chemical equilibrium concentration of CO2, which mainly depends on the prevailing temperature in the carbonator [43]. In addition to the CO2 absorption efficiency of the carbonator, the total CO2 capture efficiency of the system, Etot, is calculated according to Eq. 7.
𝐸𝑡𝑜𝑡 =
𝐹𝐶𝑂2,𝑐𝑎𝑙𝑐,𝑜𝑢𝑡 𝐹𝐶𝑂2,𝑐𝑎𝑟𝑏,𝑖𝑛 + 𝐹0 + 𝐹𝐶𝑂2,𝑆𝑅𝐹
(7)
Hereby, F0 is the molar flux of CO2 that is released by the first calcination of the fresh limestone make-up, FCO2,SRF is the CO2 that is released during the oxy-fuel combustion of SRF in the calciner. Due to the fact, that the CO2, which is formed in the oxy-fuel calciner is captured 10
nearly completely, the total CO2 capture efficiency is always higher than the CO2 absorption efficiency of the carbonator. Similar to the CO2 absorption efficiency of the carbonator, the calciner efficiency, Ecalc, is a crucial metric in the CaL process. It is defined according to Eq. 8.
𝐸𝑐𝑎𝑙𝑐 = 1 ―
𝑋𝑐𝑎𝑙𝑐
(8)
𝑋𝑐𝑎𝑟𝑏
Hereby, Xcalc is the molar carbonate content at the outlet of the calciner, and Xcarb is the molar carbonate content at the inlet of the calciner.
3
Results and Discussion
3.1
Overview
The overall goal was to demonstrate the feasibility of continuous and stable CO2 capture while using SRF as supplementary heat source. Throughout two consecutive experimental test campaigns, the pilot plant was operation more than 230 hours in coupled CFB mode under realistic CaL process conditions. This particularly implies the oxy-fuel combustion of SRF in the CFB calciner. Table 4 summarizes the range of conditions prevailing during the experimental tests. Table 4: Range of experimental test conditions. System
Carbonator
Calciner
Global
Parameter Solid inventory Superficial gas velocity Temperature CO2 conc. in primary gas CO2 absorption efficiency Solid inventory Superficial gas velocity Temperature O2 conc. in primary gas Specific sorbent circulation Specific molar make-up feed Total CO2 capture rate
Symbol Ws,carb u0,carb Tcarb yCO2,carb,in Ecarb Ws,calc u0,calc Tcalc yO2,calc,in FR/FCO2 F0/FCO2 Etot
Unit kg/m² m/s °C vol.% % kg/m² m/s °C vol.%dry mole% mole% %
Value 300 - 700 2.5 - 3.5 610 - 690 9.0 - 10.5 60 - 85 150 - 400 4.5 - 5.5 750 - 900 40 - 65 7.5 - 15 0.09 - 0.35 82 - 94
CO2 absorption efficiencies close to chemical equilibrium conditions were reached in the carbonator, while the total CO2 capture efficiency varies between 82 and 94 %. Thus, for the
11
first time, the CaL process was successfully fueled by oxy-fuel combustion of a waste derived fuel. Figure 3 shows an exemplary data plot of relevant CaL process parameters of the carbonator (Graph a) and the calciner (Graph b) for a period of six hours. During this period, the CaL process was oxy-fueled by SRF I. 100
800
a)
70
720
60 680 50
Ecarb
40 100
yi (vol.%)dry
Tc arb (°C)
760 80
Etot
Tcarb 640 900
b)
80
880
60
860
40
840
20
820 yO2,in
0
Tcalc (°C)
Ecarb , Etot (%)
90
0
1
2
3
yCO2,out 4
Tcalc 5
6
800
t (h) Figure 3: Exemplary data plot of key process parameter for carbonator (Graph a), and calciner (Graph b) while firing SRF I (F0/FCO2 ~ 0.12, FR/FCO2 ~ 11.1). Approximately 75 % of the CO2 contained in the flue gas was absorbed in the carbonator, which results in a total CO2 capture efficiency above 90 %. The solid inventory induced pressure drop was between 40 and 60 mbar in the carbonator and between 15 and 25 mbar in the calciner respectively. This corresponds to an average specific solid inventory of approximately 580 kg/m² and 280 kg/m² for carbonator and calciner, respectively. The calciner was operated within a temperature range of 820 - 850 °C, and with an inlet oxygen concentration slightly above 60 vol.%dry. During this period, the oxygen excess at the calciner outlet was between 4 and 13 vol%dry (average: 7.3 vol.%dry). This large oxygen excess was required to ensure a complete conversion of the volatile matter of SRF (xVol.,SRFI: 70 - 80 wt.%ar) at all time, while the fuel was inhomogeneous and consequently, fuel mass flow was changing considerably. In 12
addition to the high oxygen excess, the CO2 product stream at outlet of the calciner was further diluted by nitrogen required for the loop seal fluidization and the flushing of the fuel feeding system. The temperature fluctuations of the calciner were mainly caused by the inhomogeneous nature of the SRF. The carbonator was operated slightly below 680 °C. Under these circumstances, no active reactor cooling was required. This fact was caused by the lower CO2 concentration in the flue gas to be decarbonized in the carbonator. In contrast to coal-fired power plants, the CO2 concentration in the flue gas of a WtE plant is reduced by almost 30 %. Thus, less heat was released by the exothermic carbonation reaction in the carbonator. Figure 4 shows a similar data plot presenting the same data set for the carbonator (Graph a) and for the calciner (Graph b) while oxy-firing SRF II. 100
800
a)
760 80 70
720
60
Tc arb (°C)
Ecarb , Etot (%)
90
680 50
Ecarb
Etot
Tcarb
40 100
640 900
80
870
60
840
40
810
20
780 yO2,in
0
Tcalc (°C)
yi (vol.%)dry
b)
0
0.5
1
1.5
2
yCO2,out 2.5
Tcalc 3
750 3.5
t (h) Figure 4: Exemplary data plot of key process parameter for the carbonator (Graph a) and calciner (Graph b) while firing SRF II (F0/FCO2 ~ 0.22, FR/FCO2 ~ 12.2). The values for CO2 absorption efficiency and the total CO2 capture efficiency achieved during that test period are almost similar to the aforementioned results (Ecarb ~ 75 %, Etot ~ 90 %). In contrast, the overall plant operability was noteworthy different during this period. The coarse size specification of SRF II (< 50 mm) led to the accumulation of relatively large ash fractions, such as metal, glass or mineral matter in the solid looping system. This, in fact negatively 13
influences the hydrodynamics of the interconnected CFB reactors. Exemplarily, the rapid drop in the CO2 absorption efficiency within the carbonator at 0.75 h was caused by defluidization in the solid looping system due to the accumulation of coarse ash in the calciner loop seal. In order to prevent the accumulation of coarse ashes, the make-up flux was increased by almost 80 % in contrast to the operational period firing SRF I. Consequently, extraction of bottom ash was carried out more frequently. The need for a relatively large oxygen excess at the calciner outlet was also due for SRF II (xVol.,SRFII: 60 - 75 wt.%ar). 3.2
Calciner performance
The calciner efficiency (Eq. 8) is mainly depended on the calciner active space time, the prevailing CO2 concentration and the calciner temperature. The last two parameters are directly linked by the chemical equilibrium of the CaO-CO2-CaCO3 system. Accordingly, calciner temperature close to 900 °C are required to achieve a complete sorbent calcination under ideal conditions of an oxy-fired CaL calciner. At the 1 MWth CaL pilot plant, dilution of the CO2 product stream in the calciner occurred due to practical reasons such as flushing of the various pressure measurements, fluidization of the calciner loop seal and the sealing gas required in the fuel feeding system. Moreover, the oxygen excess at the calciner outlet was higher compared to previous experimental investigations of the CaL process based on hard coal or lignite as supplementary fuel at the same pilot plant [19,23,44] (see Chapter 3.1). Therefore, a sufficient sorbent calcination was feasible even at a calciner operation temperature in the range of 850 880 °C. In order to describe the calciner characteristic with regard to the chemical equilibrium conditions, the ratio, Φ, of the CO2 concentration at the outlet, yCO2,calc,out, and at the chemical equilibrium, yCO2,eq, is introduced:
𝛷=
𝑦𝐶𝑂2,𝑐𝑎𝑙𝑐,𝑜𝑢𝑡
(9)
𝑦𝐶𝑂2,𝑒𝑞
The CO2 concentration at chemical equilibrium is calculated by Eq. 10 [45]:
𝑦𝐶𝑂2,𝑒𝑞 = 4.137 ∗ 107𝑒
―
20474 𝑇𝑐𝑎𝑙𝑐
(10)
For a sufficient sorbent calcination, Φ needs to be < 1, whereas a value > 1 might lead to slow calcination or re-carbonation. These parameters are shown in Figure 5 for representative experimental conditions depended on the average calciner temperature. 14
100
2
1.5
60 1 40 0.5
20 Ecalc 0 800
Φ (-)
Ecalc (%)
80
820
840
860
880
Φ 0 900
Tcalc (°C) Figure 5: Calciner efficiency (Ecalc) and CO2 concentration ratio (Φ) depended on calciner temperature (Tcalc). Under the experimental conditions, the ratio of CO2 concentration during operation was between 0.53 and 1.75, respectively. The calciner efficiency was between 9.6 and 97 %. It can be seen that an average calciner temperature over 850 °C lead to CO2 concentration ratios below one, which results in a nearly complete sorbent calcination (Ecalc > 95 %). Once the temperature of the calciner was below 830 °C, sorbent calcination was insufficient (Ecalc < 40 %).
3.3
The Fate of Chlorine
Chlorine is introduced to the CaL process via the various waste fractions contained in the SRF. Understanding the fate of chlorine within the CaL system is crucial for an appropriate design of subsequent heat recovery systems and gas cleaning units. In this regard, any form of chlorineinduced corrosion of boiler materials is of major concern [46-48]. During oxy-fuel combustion in the CFB calciner, chlorine is released to the gas phase via HCl and partly absorbed by the solid phase. Part of solid phase bound chlorine is transported to the carbonator by the circulating sorbent. Thus, chlorine is released by the CaL process via gaseous and solid streams from the carbonator and from the calciner, respectively. The ranges of chlorine mass fraction and the total mass flow of the solid effluent streams are summarized in Table 5.
15
Table 5: Range of chlorine mass fraction and mass flow of relevant solid effluent streams from the CaL process. (BA: bottom ash, FA: fly ash). Parameter Mass flow, kg/h Mass fraction, wt.%
BACarb 2.5 - 38 0.20 - 0.32
BACalc 5.0 - 58 0.21 - 0.36
FACarb 1.0 - 21 1.61
FACalc 5.8 - 44 2.2 - 3.3
The chemical solid analysis shows that chlorine concentration is highest in the calciner fly ash, while carbonator fly ash shows a lower value in the same order of magnitude. In contrast to that, the chlorine mass fraction in bottom ashes discharged from carbonator and calciner are significantly lower. This phenomena is mainly caused by the fine Ca-particles present in the fly ash (dp,50 < 50 µm), that favor the absorption of gaseous HCl. Moreover, fly ash particles are additionally exposed to the calciner flue gases during the course of heat removal and in the bag filter. Thus, further HCl absorption is likely to occur in this process sections. Based on the chlorine introduced by SRF and the solid and gaseous effluent streams from the CaL process, a chlorine balance was established for a representative operation point (Table 6). Table 6: Chlorine balance for a representative CaL operation point2 (Tcarb: 679 °C, Tcalc: 841 °C, Etot: 83.2 %, Pth,calc: 550 kWth, FCl,SRFI: 5.42 mol/s). Parameter Cl molar flow, mol/s Cl share, %3 Cl share, %4
BACarb 0.283 0.412 0.510
BACalc 1.80 2.62 3.24
FACarb 1.61 2.34 2.89
FACalc 57.9 84.2 104
FGCalc 7.01 10.5 12.9
Sum 68.6 100 123
The fly ash of the calciner represents the major chlorine effluent from the CaL process. Moreover, it can be seen, that there is a deviation of 23 % between the chlorine feed in the SRF and the chlorine present in the effluent streams. This might be caused by several reasons, such as the fluctuation of the actual chlorine present in the SRF or the fact that the extracted solid samples represent only a minor share compared to the corresponding mass flows. Furthermore, it is likely, that HCl absorption from the calciner flue gas continues in the course of the flue gas path downstream the calciner cyclone. Due to the fact, that HCl was measured directly at the outlet of the calciner cyclone, chlorine effluent in the calciner flue gas tends to be overestimated. Knowing the concentration of HCl in the calciner flue gas, the chlorine retention rate, RCl, is calculated according to Eq. 11. This method assumes, that all gaseous chlorine in the calciner
1
Only one representative solid sample was available for the carbonator fly ash.
2
HCl in the carbonator flue gas was neglected.
3
This share refers to the sum of all chlorine in the considered effluent
4
This share refers to the chlorine in the SRF 16
off-gas is present as HCl, moreover, solid phase bound chlorine remains inert within the carbonator. Thus, there is no gaseous Cl effluent in the carbonator off-gas.
𝑅𝐶𝑙 = 1 ―
𝐹𝐻𝐶𝑙,𝑐𝑎𝑙𝑐,𝑜𝑢𝑡
(11)
𝐹𝐶𝑙,𝑆𝑅𝐹
Hereby, FHCl,FG is the molar flow of HCl in the calciner flue gas downstream of the extraction point of the recirculation gas and, FCl,fuel is the molar flow of chlorine introduced by the SRF. Figure 6 shows the average chlorine retention rate for stable operation points of one hour, each. It can be seen that the chlorine retention rate was over 82 % during all experimental test points, respectively. Furthermore, there is a tendency of an increasing chlorine retention rate along with increasing calciner temperatures. While comparing the two types of SRF, there is no distinguished difference in terms of Cl retention. As already mentioned in the course of the chlorine balance, HCl was measured directly at the outlet of the calciner cyclone. Thus, the likely further HCl absorption by Ca-rich calciner fly ash in the course of the calciner flue gas treatment path was not considered here. Therefore, chlorine retention rates tends to be even higher than those shown in Figure 6. 100 95
RCl (%)
90 85 80 SRF I SRF II
75 70 800
820
840
860
880
900
Tcalc (°C) Figure 6: Chlorine retention rate (RCl) dependent on the average calciner temperature (Tcalc).
4
Conclusions
For the first time, the CaL process has been operated while a waste derived fuel was oxy-fired in the calciner. Two types of SRF were fed to a 1 MWth pilot plant in the form of raw fluff. During more than 230 h of coupled CFB operation, CO2 was continuously captured from a flue 17
gas steam characterized by a CO2 concentration similar to typical WtE plants fueled by MSW. Under these conditions, carbonator CO2 absorption rates in the range of 80 - 85 % and total CO2 capture efficiencies over 90 % were achieved at representative CaL process conditions. It was found, that fuel quality has a significant influence on the operability of the interconnected CFB system. The pilot plant operability when using SRF I was relatively smooth and stable. When using SRF II, the enrichment of coarse ash fractions within the solid phase was the main reason that negatively affect the hydrodynamics of the interconnected CFB system. For a stable operation point, a chlorine balance reveals, that more than 80 % of the fuel-chlorine is released in the fly ash of the calciner. Due to the limestone based reactor inventory, the retention rate of chlorine was over 82 % throughout the experimental test points. It was found, that a calciner operation temperature above 860 °C is desirable in order to maximize Cl retention. When taking into account the organic waste fractions typically contained in SRF and MSW, this work successfully demonstrates the feasibility of net negative CO2 emissions by means of the CaL process at semi-industrial scale.
Acknowledgements The research leading to these results has received funding from the German Ministry of Economic Affairs and Energy based on a resolution of the German Parliament (MONIKA, FKZ: 03ET7089) and the SUEZ Group. The funding is gratefully acknowledged.
18
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Martin Haaf: Methodology, Investigation, Writing - Original Draft, Writing - Review & Editing Jens Peters: Investigation, Methodology Jochen Hilz: Investigation Antonio Unger: Resources, Supervision Jochen Ströhle: Project administration, Writing - Review & Editing, Supervision Bernd Epple: Supervision, Funding acquisition
Declaration of interests ☒ 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. ☐The authors declare the following financial interests/personal relationships which may be considered as potential competing interests:
Oxy-fuel combustion of solid recovered fuels in the calcium looping (CaL) process. Over 230 h of representative CaL operation at a 1 MWth pilot plant. Total CO2 capture rates > 90 % have been achieved. Assessment of the fate of chlorine within the calcium looping system.
24