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Modelling a two-stage countercurrent fluidized bed reactor for removing SO2from gases by active soda
Cornput. Printed
them.
Engng,
Vol.
in Great Britain
12, No.
2/3, pp. 205-208,
1988
009%1354/88 53.00 + 0.00 Pergamon Press plc
MODELLING A TWO-STAGE COUNTERCURRENT FLUIDIZED BED REACTOR FOR REMOVING SO, FROM GASES BY ACTIVE SODA L.
NEUZIL~,
F.
PROCHASKA’,
V. BUEEK’ and M.
MOCEK~
‘Department of Chemical Engineering, Prague Institute of Chemical Technology, Suchbhtarova 5, 166 28 Prague 6, Czechoslovakia. 2J. Heyrovsky Institute of Physical Chemistry and Electrochemistry, Czechoslovak Academy of Sciences, Mgchova 7, 121 38 Prague 2, Czechoslovakia (Received 8 July 1987) Abstract-The problem of high desulphurization efficiency of flue gases at high conversion in the solid phase has been investigated. Crystalline NaHCO, is fed to a reactor and by its thermal decomposition and highly reactive soda (active soda) is formed. To achieve both high process efficiency (over 95%) and high conversion of the solid (nearly 95%), we took to the two-stage countercurrent flow and to the cyclic operation of the reactor. For this purpose, the bubble assemblage model of Kato and Wen (1969) had been modified for a batch process in the solid phase including the kinetics of the chemical process taking place and the enthalpy balance. Effects of various parameters on the process have been examined by simulation. The model obtained and the results may be applied to the control of the reactor.
1. INTRODUCIlON
The AKSO process (ErdBs et al. 1968, 1982; Bare5 et al., 1970; ProchBska, 1978; Prochiska and Neuiil, 1977; Bejkk, 1983), which is a dry process for removing SO2 from flue gases, makes use of the highly reactive form of sodium carbonate, the active soda. The method is based on the thermal decomposition of NaHCO, yielding active soda: 2NaHCO,- ” I”“’ Na,CO,
+ Hz0
+ CO,,@H,)298
= 130 kJ.
(1)
This active soda reacts readily with SO, resulting in Na2S03 and Na,SO, in the presence of oxygen and preferably at elevated temperature (ProchBska, 1978; BejEek, 1983):
where X is the conversion of Na,CO,, 9 is the time since the initiation of the reaction, kR is the effective kinetic coefficient and p, is the partial pressure of SOZ. Reaction (2) is catalyzed by water vapour and for its partial pressure above 600 Pa, we have in the fluidized bed (BejEek, 1983): k, = 3.23.10-‘exp(-
1.58.10-* t),
t E( 130,2so>0c.
The reaction rate after equation (4) exerts a maximum for X = 0.25. The actual process of removing SO2 from gases by means of the active soda in a single-stage batch fluidized bed reactor is shown for dimensionless variables in Fig. 1. Desulphurization efficiency E is determined by the mol fractions SO2 in the entry and exit gaseous streams of the reactor, yi and __ y., respectively: .
Na*CO, + SO, + xi02 s(l
-x)Na,SO,
(6) + xNa,SO,
+ C02, (AH,&298= -(56.5
+x294)
The dimensionless kJ.
(2)
At a temperature below 140% mainly Na,SO, results (BejEek, 1983). In the temperature range t~<140,25O)“C, the mol fraction of the Na,SO, (xc) ratio to the mol fraction of Na,SO, (x,,) formed in the desulphurization product, is given by (Bejcek, 1983): x,/(x, + xd) = 35.9 exp( -0.025
t).
= k,p,X”*(l
- X)“*
time 0 follows from:
e GE tii,;9/n,,
(7)
where & is the mol flux of SO2 into the reactor and nb is the equivalent mols of Na,CO, present in the reactor at time 9. The relation between the variables X, 0 and E is given by
x=
E de.
(8)
(3)
The kinetics of the reaction (1) was studied by several authors (Ciborowski and Czarnota, 1963; Calistru and Ifrim, 1974). The kinetics of the reaction (2) can be expressed in the following way (Bare5 et al., 1968): dX/dS
(5)
(4) 205
Figure 1 indicates that at lower conversions, X, we obtain high values of E. While X grows, the efficiency, E, sinks below the minimal acceptable value E,. we get the so-called breakthrough of SO,. For the removing of SO, from a gaseous mixture using the AKSO process, accent is given on obtaining
L.
206
NNZIL
er al.
Enfholpy
Fig. I. Plot of X and E
for a single-stagebatch fluidized bed reactor.
vs 0
Use
bubble
Compute
as high as possible conversion in the solid phase [the maximal value (Bejtik, 1983) is X = 0.951 at a high value of E. It follows from Fig. 1 that both requirements cannot be met simultaneously in a single-stage reactor. We selected therefore a two-stage countercurrent periodically operated reactor.
balance
assemblage
time-dependent
model variables
I
terminated
2. THE MODEL
The flowsheet of the reactor assembly chosen can be seen in Fig. 2. The required mass of NaHCO,, m, is fed by the relative rate ND = tii,/& (the ratio of mol fluxes of soda and SO, in the raw gas) to the first stage where a fluidized bed is kept at the constant temperature rF, . After the cycle time interval 9,, the partially reacted contents are transferred to the 2nd stage at a conversion X,. Here it is kept at the temperature tn. The next following cycle is initiated by feeding NaHCO, to the 1st stage. Prior to reaching the reaction time 9,, the solid content is removed from stage 2 at a conversion of X,, then the solid content of the 1st stage is again transferred into the 2nd stage, etc. Gas flowing at the volume flowrate psi,.at an SOI concentration y, and temperature tgo> tF is divided into two streams during the feeding. The stream v&
5,
Yl.
*
tg*.= tF1
x=0, t.,, 1
I
YlI
tsr. =frr’
t,1i -
= t.21
YZ. \i;:
VW Y. ,t,o . V,.. y=O,twn
iI
tpe. =IF2 2
*
be \id. , %;eb YPI ,t~21
K? * tsc?. = 1F2
Fig. 2. Flow diagram of a two-stage reactor.
pALzdzz simulate
transfer
of
solid
phase
ci Yes
End
Fig. 3. Simplifiedflowchart of the computationalprogram. covers the energetic demand of the endothermic reaction (1) and it enabled to obtain the desired value of tn. Cool air at the rate r’,O is mixed with the stream Vi0 adjusting the temperature t, and decreasing the dew point of water vapour contained in the raw gas below tF. The reactor has been described by a mathematical model based on the bubble assemblage model (Kato and Wen, 1969) modified for the kinetics given by equations (4) and (5), i.e. for the reaction (2) in solid phase. The model was converted into a computer program for the simulation of the operation of a reactor. Its simplified flowchart is shown in Fig. 3. The entry data consist mainly of: geometrical dimensions of the reactor, mass of its filling mali, relative dosing rate ND, fluidized bed temperatures tF, volume flowrate psO,temperature tao and the SO, concn y,. As constant parameters we compute mainly the kinetic coefficient and the minimum fluidization velocity. The enthalpy balance is used for the determination of the distribution of
Modelling a fluid&d
a.
0
10
20
I 10
0
I 20
I
I
I
IO
20
30
I 30 0
207
(pi,, = 0) is set and the concentration yZiis computed. The computation is carried on until SO, breakthrough is attained in the 1st stage, i.e. until E, > E,,. This time 9, terminates the cycle. We compute the mean efhciency E and simulate the emptying of the 2nd stage and the transfer of the content of the 1st stage to the 2nd stage. Attaining the steady-state solid cycles is tested by comparing the values of X, in the instant of breakthrough in two successive cycles.
a(h)
dfhl
3(h)
‘)ro
bed reactor
Fig. 4. Example of a computer dependence of conversion X, and X, on time 9.
volume flowrates after Fig. 2, the SO2 concn y,, and yzr, and the superficial velocities. The computation in the following block is carried out simultaneously for both stages in chosen time intervals A9 : feeding begins at time 9 = 0. Bed height is calculated for the given time and for the given conversion X we compute the reaction rate as in equation (4). Using the bubble assemblage model, we next compute the axial concentration profiles of SO, and thus the values y,, y,, E,, E1 and or, Oz. Integrating equation (8) numerically results in new values of X, and X,. After the feeding has been terminated, new distribution of the volume flowrate
\, I
1.0
0.9
e
-.
3. DISCUSSION
To verify the influence of some input parameters on the operation of the reactor, a series of simulation calculations was carried out using the EC 1033 computer. The following input parameters were regarded as constant during the calculations: reactor diameter at the level of the grid, 0.9 m and the mean partical diameter of NaHCO,, 0.180 mm. The variable inlet parameters were as follows: SO2 concn y,~<0.008, 0.030), fluidized bed temperature t+, relative feeding rate N,~(O.85,3) and the minimal acceptable efficiency E,,o
(a)
‘,-%
*a._
‘\
lY_
-. *-..
0.6
--a, *-....-...... ....._._....
1
y.
3
2
(102)
1, f-c)
1.0
;i-1 c
\
0.9
----WV_____
...
0.6
(e)
\
. . . . . . . . .
1
&
.+! . . . . . . . . . . . .. 2
3
ND
Fig. 5. Relation between mean desulphurisation efficiency E, cycle duration g,, reaction time 9, (h) and the parameters y, (a), rF (b), m,, (c), W (d), ND (e) and E, (f).
208
L. NEUZIL et al.
tioned parameters were taken to be: y0 = 0.013, ND = 1.3, fj?= 130°C w = 5.37, m,,, = 300 kg, E, = 0.9. Ther influence of each parameter has been followed at several levels, while the remaining parameters had been kept on their basic levels. The characteristic dependence of the conversions X, and X, on time 9 is shown on Fig. 4. Time SD denotes the instant of feeding termination and 9, the reaction time of achieving X, = 0.949. As we have a shorter 9, than the cycle duration 9,, the mean efficiency in steady-state cycles obtained (_!?= 0.985) indicates that a two-stage reactor meets, in this case, the demand. To illustrate the results, Fig. 5 shows only three dependent variables E, 8, and 9s. It is obvious from the figure that mostly high desulphurization efficiencies had been achieved and that the parameter E is not too sensitive to changes of input parameters. For two cases (fF= 180°C W = 8.29), the conversions were less than 0.949 in time 9,; in these particular cases the two-stage reactor would not be sufficient. There is a maximum on the dependence of i? on ND at the value of ND--11.8 and a significant decrease of E below ND < 1. At actual conditions where both y0 and r’, vary, it is therefore an advantage to set the feeding rate in such a way, as to avoid the decrease of ND below one at maximal values of y. and r’,.. The computation time depends on the time interval chosen for a single cycle and on the duration of the cycle. For a cycle of 9, = 6.6 h and time interval AS = 30 s the computation lasted 25 s.
4. CONCLUSIONS
The results presented have shown that the twostage countercurrent fluidized bed reactor operating in cycles with the AKSO process is suitable for highly efficient purification of flue gases from SO, at a high conversion in the solid phase. In this way we can obtain a product containing a high concentration of Na,SO, or Na,SO, by purifying gases from the pulp and paper production, from the Claus process or from glass production.
REFERENCES
Bares J., J. MareEek, K. Mocek and E. Erd(is, Co&t Czech. Chem. Commun. 35. 1628 (1970). Bejfek V., AKSO process in g fluihized bed batch reactor. Ph.D. Thesis, Prague Institute of Chemical Technology_. (1983). Calistru C. and L. Ifrim, Bufetinal lnstitutului Politehnic Iasi 2% 45 (1974). Ciborowski J. and T. Czamota, Przem. them. 42, 313 (1963). Erdos E., J. Bare& J. MareEek and K. Mocek, Cs. Pat. No. 129889 f 1968). Erdijs E., K. Mbcek, J. Mar&ek and I. Bares, Property of dry solid sodium carbonate. Cr. Discovery No. 26 (1982). Kato K. and C. Y. Wen, Chem. Engng Sci. 24,135l (1969). Prochhska F., Modelling fluidized bed reactor. Ph. D. Thesis, Prague Institute of Chemical Technology (1978). Prochaska F. and L. Netiil, Removing sulfur dioxide from gaseous mixtures in a fluidized bad formed of active soldium carbonate. Proc. 3rd Conf. Appl. Chem., Veszprem (1977).
them.
Engng,
Vol.
in Great Britain
12, No.
2/3, pp. 205-208,
1988
009%1354/88 53.00 + 0.00 Pergamon Press plc
MODELLING A TWO-STAGE COUNTERCURRENT FLUIDIZED BED REACTOR FOR REMOVING SO, FROM GASES BY ACTIVE SODA L.
NEUZIL~,
F.
PROCHASKA’,
V. BUEEK’ and M.
MOCEK~
‘Department of Chemical Engineering, Prague Institute of Chemical Technology, Suchbhtarova 5, 166 28 Prague 6, Czechoslovakia. 2J. Heyrovsky Institute of Physical Chemistry and Electrochemistry, Czechoslovak Academy of Sciences, Mgchova 7, 121 38 Prague 2, Czechoslovakia (Received 8 July 1987) Abstract-The problem of high desulphurization efficiency of flue gases at high conversion in the solid phase has been investigated. Crystalline NaHCO, is fed to a reactor and by its thermal decomposition and highly reactive soda (active soda) is formed. To achieve both high process efficiency (over 95%) and high conversion of the solid (nearly 95%), we took to the two-stage countercurrent flow and to the cyclic operation of the reactor. For this purpose, the bubble assemblage model of Kato and Wen (1969) had been modified for a batch process in the solid phase including the kinetics of the chemical process taking place and the enthalpy balance. Effects of various parameters on the process have been examined by simulation. The model obtained and the results may be applied to the control of the reactor.
1. INTRODUCIlON
The AKSO process (ErdBs et al. 1968, 1982; Bare5 et al., 1970; ProchBska, 1978; Prochiska and Neuiil, 1977; Bejkk, 1983), which is a dry process for removing SO2 from flue gases, makes use of the highly reactive form of sodium carbonate, the active soda. The method is based on the thermal decomposition of NaHCO, yielding active soda: 2NaHCO,- ” I”“’ Na,CO,
+ Hz0
+ CO,,@H,)298
= 130 kJ.
(1)
This active soda reacts readily with SO, resulting in Na2S03 and Na,SO, in the presence of oxygen and preferably at elevated temperature (ProchBska, 1978; BejEek, 1983):
where X is the conversion of Na,CO,, 9 is the time since the initiation of the reaction, kR is the effective kinetic coefficient and p, is the partial pressure of SOZ. Reaction (2) is catalyzed by water vapour and for its partial pressure above 600 Pa, we have in the fluidized bed (BejEek, 1983): k, = 3.23.10-‘exp(-
1.58.10-* t),
t E( 130,2so>0c.
The reaction rate after equation (4) exerts a maximum for X = 0.25. The actual process of removing SO2 from gases by means of the active soda in a single-stage batch fluidized bed reactor is shown for dimensionless variables in Fig. 1. Desulphurization efficiency E is determined by the mol fractions SO2 in the entry and exit gaseous streams of the reactor, yi and __ y., respectively: .
Na*CO, + SO, + xi02 s(l
-x)Na,SO,
(6) + xNa,SO,
+ C02, (AH,&298= -(56.5
+x294)
The dimensionless kJ.
(2)
At a temperature below 140% mainly Na,SO, results (BejEek, 1983). In the temperature range t~<140,25O)“C, the mol fraction of the Na,SO, (xc) ratio to the mol fraction of Na,SO, (x,,) formed in the desulphurization product, is given by (Bejcek, 1983): x,/(x, + xd) = 35.9 exp( -0.025
t).
= k,p,X”*(l
- X)“*
time 0 follows from:
e GE tii,;9/n,,
(7)
where & is the mol flux of SO2 into the reactor and nb is the equivalent mols of Na,CO, present in the reactor at time 9. The relation between the variables X, 0 and E is given by
x=
E de.
(8)
(3)
The kinetics of the reaction (1) was studied by several authors (Ciborowski and Czarnota, 1963; Calistru and Ifrim, 1974). The kinetics of the reaction (2) can be expressed in the following way (Bare5 et al., 1968): dX/dS
(5)
(4) 205
Figure 1 indicates that at lower conversions, X, we obtain high values of E. While X grows, the efficiency, E, sinks below the minimal acceptable value E,. we get the so-called breakthrough of SO,. For the removing of SO, from a gaseous mixture using the AKSO process, accent is given on obtaining
L.
206
NNZIL
er al.
Enfholpy
Fig. I. Plot of X and E
for a single-stagebatch fluidized bed reactor.
vs 0
Use
bubble
Compute
as high as possible conversion in the solid phase [the maximal value (Bejtik, 1983) is X = 0.951 at a high value of E. It follows from Fig. 1 that both requirements cannot be met simultaneously in a single-stage reactor. We selected therefore a two-stage countercurrent periodically operated reactor.
balance
assemblage
time-dependent
model variables
I
terminated
2. THE MODEL
The flowsheet of the reactor assembly chosen can be seen in Fig. 2. The required mass of NaHCO,, m, is fed by the relative rate ND = tii,/& (the ratio of mol fluxes of soda and SO, in the raw gas) to the first stage where a fluidized bed is kept at the constant temperature rF, . After the cycle time interval 9,, the partially reacted contents are transferred to the 2nd stage at a conversion X,. Here it is kept at the temperature tn. The next following cycle is initiated by feeding NaHCO, to the 1st stage. Prior to reaching the reaction time 9,, the solid content is removed from stage 2 at a conversion of X,, then the solid content of the 1st stage is again transferred into the 2nd stage, etc. Gas flowing at the volume flowrate psi,.at an SOI concentration y, and temperature tgo> tF is divided into two streams during the feeding. The stream v&
5,
Yl.
*
tg*.= tF1
x=0, t.,, 1
I
YlI
tsr. =frr’
t,1i -
= t.21
YZ. \i;:
VW Y. ,t,o . V,.. y=O,twn
iI
tpe. =IF2 2
*
be \id. , %;eb YPI ,t~21
K? * tsc?. = 1F2
Fig. 2. Flow diagram of a two-stage reactor.
pALzdzz simulate
transfer
of
solid
phase
ci Yes
End
Fig. 3. Simplifiedflowchart of the computationalprogram. covers the energetic demand of the endothermic reaction (1) and it enabled to obtain the desired value of tn. Cool air at the rate r’,O is mixed with the stream Vi0 adjusting the temperature t, and decreasing the dew point of water vapour contained in the raw gas below tF. The reactor has been described by a mathematical model based on the bubble assemblage model (Kato and Wen, 1969) modified for the kinetics given by equations (4) and (5), i.e. for the reaction (2) in solid phase. The model was converted into a computer program for the simulation of the operation of a reactor. Its simplified flowchart is shown in Fig. 3. The entry data consist mainly of: geometrical dimensions of the reactor, mass of its filling mali, relative dosing rate ND, fluidized bed temperatures tF, volume flowrate psO,temperature tao and the SO, concn y,. As constant parameters we compute mainly the kinetic coefficient and the minimum fluidization velocity. The enthalpy balance is used for the determination of the distribution of
Modelling a fluid&d
a.
0
10
20
I 10
0
I 20
I
I
I
IO
20
30
I 30 0
207
(pi,, = 0) is set and the concentration yZiis computed. The computation is carried on until SO, breakthrough is attained in the 1st stage, i.e. until E, > E,,. This time 9, terminates the cycle. We compute the mean efhciency E and simulate the emptying of the 2nd stage and the transfer of the content of the 1st stage to the 2nd stage. Attaining the steady-state solid cycles is tested by comparing the values of X, in the instant of breakthrough in two successive cycles.
a(h)
dfhl
3(h)
‘)ro
bed reactor
Fig. 4. Example of a computer dependence of conversion X, and X, on time 9.
volume flowrates after Fig. 2, the SO2 concn y,, and yzr, and the superficial velocities. The computation in the following block is carried out simultaneously for both stages in chosen time intervals A9 : feeding begins at time 9 = 0. Bed height is calculated for the given time and for the given conversion X we compute the reaction rate as in equation (4). Using the bubble assemblage model, we next compute the axial concentration profiles of SO, and thus the values y,, y,, E,, E1 and or, Oz. Integrating equation (8) numerically results in new values of X, and X,. After the feeding has been terminated, new distribution of the volume flowrate
\, I
1.0
0.9
e
-.
3. DISCUSSION
To verify the influence of some input parameters on the operation of the reactor, a series of simulation calculations was carried out using the EC 1033 computer. The following input parameters were regarded as constant during the calculations: reactor diameter at the level of the grid, 0.9 m and the mean partical diameter of NaHCO,, 0.180 mm. The variable inlet parameters were as follows: SO2 concn y,~<0.008, 0.030), fluidized bed temperature t+
(a)
‘,-%
*a._
‘\
lY_
-. *-..
0.6
--a, *-....-...... ....._._....
1
y.
3
2
(102)
1, f-c)
1.0
;i-1 c
\
0.9
----WV_____
...
0.6
(e)
\
. . . . . . . . .
1
&
.+! . . . . . . . . . . . .. 2
3
ND
Fig. 5. Relation between mean desulphurisation efficiency E, cycle duration g,, reaction time 9, (h) and the parameters y, (a), rF (b), m,, (c), W (d), ND (e) and E, (f).
208
L. NEUZIL et al.
tioned parameters were taken to be: y0 = 0.013, ND = 1.3, fj?= 130°C w = 5.37, m,,, = 300 kg, E, = 0.9. Ther influence of each parameter has been followed at several levels, while the remaining parameters had been kept on their basic levels. The characteristic dependence of the conversions X, and X, on time 9 is shown on Fig. 4. Time SD denotes the instant of feeding termination and 9, the reaction time of achieving X, = 0.949. As we have a shorter 9, than the cycle duration 9,, the mean efficiency in steady-state cycles obtained (_!?= 0.985) indicates that a two-stage reactor meets, in this case, the demand. To illustrate the results, Fig. 5 shows only three dependent variables E, 8, and 9s. It is obvious from the figure that mostly high desulphurization efficiencies had been achieved and that the parameter E is not too sensitive to changes of input parameters. For two cases (fF= 180°C W = 8.29), the conversions were less than 0.949 in time 9,; in these particular cases the two-stage reactor would not be sufficient. There is a maximum on the dependence of i? on ND at the value of ND--11.8 and a significant decrease of E below ND < 1. At actual conditions where both y0 and r’, vary, it is therefore an advantage to set the feeding rate in such a way, as to avoid the decrease of ND below one at maximal values of y. and r’,.. The computation time depends on the time interval chosen for a single cycle and on the duration of the cycle. For a cycle of 9, = 6.6 h and time interval AS = 30 s the computation lasted 25 s.
4. CONCLUSIONS
The results presented have shown that the twostage countercurrent fluidized bed reactor operating in cycles with the AKSO process is suitable for highly efficient purification of flue gases from SO, at a high conversion in the solid phase. In this way we can obtain a product containing a high concentration of Na,SO, or Na,SO, by purifying gases from the pulp and paper production, from the Claus process or from glass production.
REFERENCES
Bares J., J. MareEek, K. Mocek and E. Erd(is, Co&t Czech. Chem. Commun. 35. 1628 (1970). Bejfek V., AKSO process in g fluihized bed batch reactor. Ph.D. Thesis, Prague Institute of Chemical Technology_. (1983). Calistru C. and L. Ifrim, Bufetinal lnstitutului Politehnic Iasi 2% 45 (1974). Ciborowski J. and T. Czamota, Przem. them. 42, 313 (1963). Erdos E., J. Bare& J. MareEek and K. Mocek, Cs. Pat. No. 129889 f 1968). Erdijs E., K. Mbcek, J. Mar&ek and I. Bares, Property of dry solid sodium carbonate. Cr. Discovery No. 26 (1982). Kato K. and C. Y. Wen, Chem. Engng Sci. 24,135l (1969). Prochhska F., Modelling fluidized bed reactor. Ph. D. Thesis, Prague Institute of Chemical Technology (1978). Prochaska F. and L. Netiil, Removing sulfur dioxide from gaseous mixtures in a fluidized bad formed of active soldium carbonate. Proc. 3rd Conf. Appl. Chem., Veszprem (1977).