カルシウムルーピングプロセスによる CO2 と N2O の同時排出低減
Simultaneous reduction of CO2 and N2O emissions by Calcium-Looping Process
新潟大学工学部化学システム工学プログラム
教授 清水 忠明
Tadaaki Shimizu
Professor, Program of Chemistry and Chemical Engineering, Niigata University
Introduction
Calcium Looping (CaL) process, which uses CaO particles as sorbent, has beendeveloped as a post-combustion CO2 capture process as portrayed in Fig. 11). In thisprocess, dual-fluidized bed solid circulation system is usually employed. Carbon
dioxide in flue gas from an air-blown combustor is captured by CaO in the
carbonator at about 873 K. Decomposition of CaCO3 to CaO and CO2 is conducted in
the regenerator at about 1223 K. The heat to decompose CaCO3 in the regenerator
is supplied by burning fuel such as coal with pure O2 diluted by recirculated CO2 sothat the flue gas consists ideally only of CO2 and H2O. The CaL process is regarded
as an energy-efficient and low-cost post-combustion CO2 capture process1)-6).
Therefore, CaL process has been extensively investigated using mainly
dual-fluidized bed systems to demonstrate CO2 capture7)-20).
CO 2 , H 2 O

CO 2 -free gas
CaO
CaO+CO 2→ CaCO 3

CaCO 3

CaCO 3 →CaO+CO 2

Flue gas(CO 2 , N 2 O, NO x , O 2 )Fuel O 2

Carbonator

Regenerator

Fig. 1

CO 2

Concept of CaL process

In addition to CO2 capture, behavior of N2O in the carbonator of CaL process isconcern. N2O is a greenhouse gas with high global warming potential (GWP) of
29821). Among air-blown coal combustors, circulating fluidized bed combustors(CFBCs) are known to release flue gas with relatively high N2O concentrations ofabout 100–200 ppm (equivalent to 3–6% CO2) during coal combustion 22)-25). Since
the sorbent (CaO) particles are known to decompose N2O at elevatedtemperatures26)-33), CaL process is expected to work as a process for N2O
decomposition as well as CO2 capture, both of which contribute to the reduction ofgreenhouse gases.
The objective of this work is to clarify the reaction pathways of CO2 and N2O inthe carbonator of CaL process. This work consists of three studies, 1)bubbling
fluidized bed study to observe the fate of N2O in the presence of CaO under－ 173 －

carbonator conditions, 2)dual-fluidized bed solid circulation system study todemonstrate N2O decomposition and CO2 capture, and 3)two-stage fluidized bed
study to simulate solid circulation in the dual-fluidized bed solid circulation system.
The last one, the two-stage system study, was designed to observe the behavior ofN2O decomposition and CO2 capture under well-defined solid feed conditions; it is
not easy to determine the solid circulation rate in the dual-fluidized bed system
operated at high temperatures. Thus in the two-stage system study, apreviously-calibrated screw feeder was employed to feed the sorbent at a desired
feed rate.
Experimental
1. Bubbling fluidized bed experiments
Fig. 2 presents an illustration of a schematic diagram of a laboratory-scale
bubbling fluidized bed system used for the present work. This fluidized bed wasmade of a quartz tube that had 26 mm inner diameter and total height of 1 m. A
quartz sintered plate gas distributor was installed to support the solid sample. The
solid sample was packed onto the distributor with a static bed height of 7 cm. The
bed temperature was measured using a thermocouple (JIS Type-K). An electricheater heated the reactor to a fixed temperature of 873 K. A gas mixture of N2O andO2 diluted by N2 was fed from the bottom. The superficial gas velocity was fixed at
17 cm/s at 873 K, i.e. the gas flow rate was 1.7 NL/min. The respective
concentrations of N2O and O2 were fixed at 480ppm and 4%.
Calcined limestone was employed as the bed material. The composition (wt.%) of
raw limestone included CaCO3 96.9, MgCO3 1.4, SiO2 0.6, Al2O3 0.8, and Fe2O3 0.3.
The raw limestone particle size was 0.35–0.42 mm. The limestone was calcined inN2 stream at 1173 K.
Char
FilterN 2 (for dilution andflow rate increase)TR
CO 2 & CO
Analyzer

Quartzreactor
CaO/sand
bed

HeaterNOx
Analyzer

Gas bags
O2
NO
N2 O
CO
CO 2
N2

Fig. 2

Schematic diagram of a laboratory-scale bubbling fluidized bed system

2. Dual-fluidized bed experiments
Fig. 3 illustrates a schematic diagram of dual-fluidized bed system employed forthe present work. This dual-fluidized bed system consisted of a bubbling fluidized－ 174 －

bed carbonator and a fast fluidized bed regenerator. The detail of the apparatus isgiven elsewhere34)35). The carbonator was 9.3 cm in inner diameter and the bedheight was fixed at 0.30 m by an overflow tube. The particles from the overflow tube
were recirculated back to the bottom of the regenerator through a loopseal. The
inner diameter and the height of the regenerator were 2.2 cm and 1.93 m,
respectively. The particles were transported to the top of the regenerator by the
upward gas stream, separated from the flue gas by a cyclone, and introduced into
the carbonator. Gas mixture simulating flue gas from air-blown combustor consistedof CO2, N2O, O2, and N2 was fed to the carbonator. The concentration of CO2 in the
carbonator fluidizing gas was fixed at 15%. The concentration of N2O was fixed at
240ppm. The O2 concentration in the carbonator fluidizing gas was varied from 4%
to 16.8%. The temperature in the carbonator was fixed at 873 K. The superficial gas
velocity was 0.16 m/s at this temperature. The regenerator was fluidized by air
without fuel feeding so that material balance of CO2 could be evaluated. The
superficial air velocity above the secondary gas inlet was fixed at 2.75 m/s (corrected
at 1223 K). Gas leakage between two reactors was evaluated by feeding He and Ne
as tracer gases to the regenerator and the carbonator, respectively. The gas leakage
was found to be negligible.
The same limestone with different size from the bubbling fluidized bed
experiments was employed as the bed material. The Sauter mean diameter of the
raw limestone was 0.268 mm. The maximum size was 0.42 mm.
Regenerator fluegas to analyzers
CycloneSolids(tracer)

CaO
HopperCarbonator fluegas to analyzers

Regenerator(Fast FB)Δ P U bydifference

Carbonator(Bubbling FB)

ΔP RSecondary gas

Simulated flue gas(CO 2 , N 2 O, NO, O 2 , N 2 )T
Thermocouples to T
measure downwardTvelocity of solids

Stand
pipe
CaO/
CaCO 3

ΔP B

Loopseal
Loopsealgas

Bottom fluidizing gas
Primary gas

Fig. 3 Schematic diagram of dual-fluidized bed solid circulating system
3. Two-stage bubbling fluidized bed experiments
Fig. 4 illustrates a schematic diagram of two-stage fluidized bed system
employed for the present work. This two-stage fluidized bed system consisted of two
bubbling fluidized bed reactors of the same size connected in series. The inner－ 175 －

diameter and total height were 3.7 cm and 44 cm, respectively. An overflow tube
was installed at 10 cm above the gas distributor. The first-stage fluidized bed
reactor was used as a carbonator to which calcined limestone and simulated flue gas
were fed. As the simulated flue gas, a mixture of CO2, O2, and N2O diluted by N2
was fed. The respective concentrations of N2O and O2 were fixed at 1000ppm and
4%. Superficial gas velocity was fixed at 0.10 m/s.
Calcined limestone particles were continuously fed from a hopper though ascrew feeder to the top of the first-stage fluidized bed continuously. After partial
carbonation, the solids withdrawn through an overflow tube of the first-stage
fluidized bed was transported via a loopseal to the second-stage fluidized bed. The
particles in loopseal were fluidized by N2. The second-stage fluidized bed was used
as a regenerator, into which nitrogen was fed as fluidizing gas to calcine CaCO3. The
calcined limestone particles withdrawn through an overflow tube of thesecond-stage fluidized bed was stored in a container purging the gas with inert gas(N2). The solids stored in the container was then fed to the CaO hopper again, thus
experiment was continued with solid sample using increased number of cycle. The
condition of the regenerator was designed to establish the material balance of CO2
capture in the carbonator and release of CO2 in the regenerator, thus CO2-free gas
was fed to the regenerator.
For the calculation of the material balance, the split of fluidizing gas fed to theloopseal to the first stage and the second stage should be known. For the evaluationof the split, tracer gas (He) was fed to the loopseal fluidizing gas. The concentrationof He in the gas samples was measured by gas chromatography. The tracer gas was
detected in the gas samples from both the first-stage reactor and the second-stage
reactor. This result indicates that the gas fed to the loopseal moved upward not only
through the standpipe connecting the carbonator and loopseal but also the upward
flow section of the loopseal connecting the regenerator and the loopseal. When the
direct gas leakage occurred from the carbonator to the regenerator, He could notreach the carbonator. The fact that He was detected in the gas from carbonatorindicated that upward flow of gas was formed in the downcomer of the loopseal.
The bed temperature was measured by use of thermocouples and controlled byuse of electric heaters so that suitable temperatures for carbonation (873 K) andcalcination (1223 K) were attained in the carbonator and the regenerator,
respectively. The pressure drop between the bottom of the dense bed and thefreeboard was measured to measure the solid hold-up in the reactors.
Calcined limestone sample after repeated use in the dual-fluidized bed system
was employed. The solid sample was fed through a screw feeder to the top of thefirst-stage reactor. The relationship between the rotating speed of the screw and thefeed rate of solids by bulk volume was calibrated in advance. Also the descending
rate of the solid’s surface in the hopper was measured intermittently, thus the bulk
volume flow rate of the solids was calculated. By measuring the bulk density of thesolids before packing into the solid hopper, the mass flow rate of the solids wasestimated by multiplying bulk volume flow rate and the bulk density.

－ 176 －

CaO hopper

Sorbent repeatedly used

Screw feeder
M MotorTRto gas analyzers

1st-stage FB(Carbonator)

ΔP 1
+

-

ΔPF

-

+

ΔPL

Simulated flue gas(CO 2 , CO, NO, N 2 O , N 2 )+TR

LoopsealN2
He (tracer)

to gas analyzers
2nd-stage FB(Regenerator)

-

ΔP2
+

N 2 for purge
N2
Container

Fig. 4

Schematic diagram of two-stage fluidized bed system

The detail of the gas analysis is given elsewhere21)22). The CO2 and CO
concentrations in the flue gas from the reactors were measured using NDIR
absorption. Concentrations of NOx were continuously analyzed using chemical
luminescence NOx analyzers. Flue gas samples were intermittently stored in gas
bags. Concentrations of N2O, He, and Ne in the gas samples were measured using
gas chromatography with a thermal conductivity detector.
Results and Discussion
1. Bubbling fluidized bed experiments
The decomposition of N2O was measured. When inert silica sand was employed
as the bed material, the unreacted fraction of N2O was about 98%, i.e., thermal
decomposition of N2O in the present fluidized bed apparatus was negligible. When
CaO was employed, N2O concentration in the produced gas was lower than thedetection limit (2ppm), i.e. more than 99% of N2O was decomposed in the CaO bed.
This decomposition is attributable to the catalytic activity to decompose N2O atelevated temperatures 26)-33). Thus calcined limestone can work as a catalyst for N2O
decomposition under a temperature condition of carbonator.
2. Dual-fluidized bed experiments
Fig. 5 shows a result of CO2 capture in the carbonator and CO2 release from theregenerator of dual-fluidized bed. Immediately after the start of CO2 feed to thecarbonator, the outlet flow rate of CO2 from the carbonator decreased remarkably,i.e., most of fed CO2 was captured by CaO particles in the carbonator. With the－ 177 －

accumulation of captured CO2 in the system, the rate of CO2 release in theregenerator increased and finally the release rate of CO2 reached the feed rate of
CO2 to the carbonator at about 1 hour after the start, i.e., the material balance of
CO2 was closed.

CO2 flow [mmol/s]3(a)Start of CO feed
to Carb.

O2 in Carb. feed gas = 4.0%

2

2

1
CO2
CO2 release
release fromfrom Reg.
Reg.
Carb.
Carb.CO2
CO2 feed
feed

0
14600

CO2 flow [mmol/s]

3

15800

17000

18200

19400
Time [s]

20600

21800

23000

O2 in Carb. feed gas =16.8%(b)Start of CO2 feed to Carb.

2

1

CO2
CO2flow
flowatatCarb.
Carb.outlet
outletCO2
CO2 feed
feed rate
rate toto Carb.
Carb.

CO2 flow
flow at
at Reg.
Reg. outlet
outletCO2
0
1000

Fig. 5

2200

3400

4600
Time[s]

5800

7000

8200

CO2 capture in carbonator and CO2 release from regenerator ofdual-fluidized bed

Fig. 6 shows the decomposition of N2O in the carbonator of dual-fluidized bed. As
expected from the results of bubbling bed experiments, the decomposition of N2O
was observed in the carbonator. Though the contact time (=solid volume/gas volume
flow rate) of the carbonator of the dual-fluidized bed was longer than that of the
bubbling bed apparatus, the conversion of N2O observed for the carbonator of thedual-fluidized bed was lower than that of the bubbling fluidized bed. This is partly
attributable to the decrease in solid activity after repeating carbonation/calcination
cycles. For the bubbling bed experiments, fresh CaO particles after calcination of
raw limestone were used, whereas the CaO particles used during dual-fluidized bed
experiments had experienced a number of carbonation/calcination cycles. It is
known that the CaO particles lose activity for CO2 capture after repeating
carbonation/calcination cycles because of sintering of grains1) 36) 37). The reduced CO2
capture capacity is attributable to the reduction of surface area through sintering ofgrains or plugging of small pores 38) 39) 40). Thus the activity to decompose N2O is also
considered to be reduced after repeating carbonation and calcination cycles.

－ 178 －

N2O in Carb. flue gas [ppm]

300
250

O2 in Carb. feed gas =16.8%200

N2O in Carb. flue gas
N2O

150

N2O in Carb. feed gas
N2O

100
50
0
1000

Fig. 6

2200

3400

4600
Time [s]

5800

7000

8200

Decomposition of N2O in carbonator of dual-fluidized bed

3. Two-stage bubbling fluidized bed experiments
Fig. 7 shows the flow rates of CO2 from the carbonator and the regenerator fordifferent CaO feed rates. When CaO was fed, CO2 was captured in the first-stage
carbonator and the decomposition of CaCO3 took place in the second-stage
regenerator. Comparing the results obtained using the dual-fluidized bed system(Fig. 5), less CO2 capture was observed. This difference is attributable to the lower
bed height of the two-stage system (10 cm) compared with that in the dual-fluidized
bed system (30 cm). For the analysis of the results, the present dual-fluidized bed
system captured too much CO2 to analyze the results; i.e. when the conversion of
the reactant exceeds about 98%, the outlet concentration of the reactant is too low todetermine the reaction rate from the ratio of the concentration at the inlet to that at
the outlet. Thus conversion of the solid reactant was determined by the results ofthe two-stage fluidized bed system because both the inlet concentration and theoutlet concentration could be determined with high accuracy.

CO2 flow rate [mmol/s]

0.5

CO2
fromCarb.
Carb.
CO2 from
CaO feed = 0.105 mol/min

0

Carb.
CO2
feed rate
Carb. CO
2 feed rate.
0.192

0

0.4
0.3
0.2
0.1
0
16000

Fig. 7

CO2
fromReg.
Reg.
CO2 from
0.162

17000

18000

19000

20000
Time[s]

21000

22000

23000

Flow rate of CO2 from carbonator and regenerator compared with CO2 feed
rate to carbonator observed using two-stage fluidized bed system

Fig. 8 shows the influence of CaO feed rate on CO2 capture rate in the carbonatorand conversion of CaO to CaCO3 observed using two-stage fluidized bed system,
where the conversion was calculated from the CaO feed rate and CO2 capture rate.
With increasing CaO feed rate, CO2 capture rate in the carbonator increased, but－ 179 －

0.010

0.1

0.008

0.08

0.006

0.06

0.004

0.04

0.000

Fig. 8Capture rate

0.002

0.02

Conversion of CaO
0.00

0.05 0.10 0.15 0.20
CaO feed rate [mol/min]

0.25

Conv. of CaO to CaCO3 [-]CO2 capture rate [mol/min]the conversion of CaO decreased.

0

Influence of CaO feed rate on CO2 capture rate in carbonator and conversion
of CaO to CaCO3 observed using two-stage fluidized bed system

Fig. 9 shows N2O concentration normalized by N2 concentration in the gas fed tocarbonator and the flue gas from the carbonator. Because of the decrease in total
gas flow rate after CO2 capture, direct comparison of the concentration of N2O in the
fed gas with that at the outlet is meaningless, N2 was employed for the inert tracer
for the analysis of the conversion of N2O. About 15% of fed N2O was found to be
decomposed in the carbonator. The decomposition of N2O is attributable to thecatalytic activity of CaO. The conversion of N2O observed in the present two-stage
fluidized bed system was lower than that of the present dual-fluidized bed system(Fig. 6). This difference is attributable to the difference in the bed height as
discussed above.
Carb.N2O/N2 ratio [mmol/mol]

1.0
CaO feed = 0.105 mol/min

0.162

0

0.192

0

0.5

0.0
16000

Fig. 9

17000

18000

Carb. outlet

Carb. inlet

19000

21000

20000
Time[s]

22000

23000

N2O decomposition in carbonator of two-stage fluidized bed system

Conclusion
The present work revealed that the sorbent of CaL process can work as acatalyst to decompose N2O under a temperature condition of the carbonator of CaL
process. Dual fluidized bed experiments revealed that a part of N2O was－ 180 －

decomposed in the carbonator during continuous solid circulation with carbonationof CaO and calcination of produced CaCO3. The two-stage fluidized bed experiments
enabled the continuous model experiments of CaL process under precise solid feed
conditions to obtain kinetic data of CO2 capture and N2O decomposition. CaL can bea multi-functional process for not only CO2 separation but also for N2O
decomposition.
Acknowledgement
T. Shimizu expresses his appreciation for the financial support from JFE 21stCentury Foundation.
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