Scale-up of Circulating Fluidized Bed Combustion - Energy & Fuels

1286

Energy & Fuels 2000, 14, 1286-1292

Scale-up of Circulating Fluidized Bed Combustion Bo Leckner* Department of Energy Conversion, Chalmers Technical University, 41296 Go¨ teborg, Sweden

Joachim Werther† Chemical Engineering I, Technical University Hamburg-Harburg, Denickestrasse 15, D21073 Hamburg, Germany Received May 23, 2000. Revised Manuscript Received August 9, 2000

Tests on large circulating fluidized bed boilers are extremely expensive and computer modeling is not yet reliable. The paper tries to answer the question: Can tests on laboratory-scale equipment be scaled to satisfy the need for information? Criteria for combustion scaling are defined in the form of two Damko¨hler numbers: one for processes in the vertical direction and one for the horizontal direction. Examples are given to show the outcome of combustion modeling. It is concluded that it is possible to reasonably well model the combustion process in the vertical direction by providing sufficient residence time in the riser, but in a laboratory device it is not possible to model the horizontal combustion processes, except in special cases.

Introduction Tests on large circulating fluidized bed (CFB) boilers are extremely expensive and computer modeling is not yet reliable. Therefore, the following question must be answered: are tests on scale models the solution? In fact, it is a common practice to make tests in the laboratoryson laboratory scale equipment. The results from such tests are presented as if they were generally valid, but active scale modeling related to large-scale combustors is usually not carried out. This is the background to the discussion below regarding the question: to what extent are combustion results from laboratory fluidized bed reactors valid for commercial scale combustors? Scope of the Problem. There are two processes that have to be considered, fluid dynamics, and combustion. Fluid Dynamic Scaling. When the gas velocity is increased in a fluidized bed, the bubbles increase in size. In laboratory scale beds, operated with velocities up to those of CFB combustors and with normal bed materials, an influence from the walls of the fluidization vessel cannot be avoided as the bubbles grow. This leads to slugging and transition to the turbulent regime of fluidization, as described in numerous publications. In wide combustors, such as in commercial boilers, the walls do not interact: the bubbles can develop in size until they break up (Group A powders) or until they reach the size of the bed height (Group B powders). As a consequence, the fluidization conditions in the bottom region of a riser are quite different in a narrow laboratory riser and in a boiler. This has not caused much * To whom correspondence should be addressed. Phone: +46 31 7721431. Fax: +46 31 77233592. E-mail: energy.conversion@entek. chalmers.se. † Phone: +49 40 42878 3039. Fax: +49 40 42878 2678. E-mail: [email protected].

discussion, since most studies have avoided the bottom part of a riser and have focused attention on the central or upper part, where the conditions are qualitatively similar in narrow and wide risers. In a combustor, however, the bottom region is important, since much of the conversion takes place there, and this is the reason why particular attention should be paid to this part of the riser. Figure 11 gives an example of risers studied, taken from publications before 1995. The figure illustrates the small aspect ratio of commercial scale combustors (upper part of the diagram) compared to the large aspect ratio of the tall and narrow CFB laboratory risers. The black part of each bar indicates where in the risers the experiments were carried out (in all cases, far from the bottom). Fluid dynamic scaling is the solution to the problem described. If the scaling is properly carried out the bubbles in the small laboratory scale apparatus should be correspondingly small and equivalent to the larger bubbles in the large-scale device; the fluid dynamics should be scaled. Criteria for such scaling have been derived by, among others, Glicksman et al.2 They have compared the behavior of scale models and large plants and found a reasonable agreement, at least with respect to particular characteristics, such as frequencies of pressure fluctuations. Another study of this kind3 compared pressure measurements in a 12MW boiler with corresponding data from a cold test rig, a 1/9 scale (1) Johnsson, F.; Zhang, W.; Leckner, B. Proceedings of the 2nd International Conference on Multiphase Flow; Serizawa, A., Fukano, T., Battaille, J., Ed.; Japan Society of Multiphase Flow: Kyoto, 1995; Vol. 3, FB1-25. (2) Glicksman, L. R.; Hyre, M. R.; Farrell, P. A. Int. J. Multiphase Flow 1994, 20, Suppl. 331-386. (3) Johnsson, F.; Vrager, A.; Leckner, B. Proceedings of the 15th International Conference on FBC; Reuther, R. B., Ed; ASME: New York, 1999; Paper FBC99-0047.

10.1021/ef0001078 CCC: $19.00 © 2000 American Chemical Society Published on Web 10/27/2000

Fluidized Bed Combustion

Energy & Fuels, Vol. 14, No. 6, 2000 1287

mathematical modeling. Therefore, simplified relationships will be formulated here to catch only the most essential features of the combustion processes in a CFB furnace. The Damko¨hler number,6 Da, is proposed as a criterion to determine the combustion scaling. In general terms this number can be defined as

Da ) transport time/reaction time

Figure 1. Example of shapes (aspect ratio) and location of measurements in risers studied in published reports.1 The reference numbers in the figure are given in ref 1. (H is the total height of the riser, D its equivalent (hydraulic) diameter).

model of the same boiler. Both the vertical suspension density profile and the frequency of pressure fluctuations were well represented by the scale model. Only minor deviations were observed, which may have originated from deviations in the properties of the particles used from those prescribed by the scaling criteria (there is a fundamental difficulty in finding particles with exactly the prescribed size, density and sphericity). However, the aim of the present work is to study combustion. Then, other criteria become important, and these criteria may not be at all compatible with those of fluid dynamic scaling. Combustion Scaling. A great step toward combustion scaling is taken already if important parameters can be kept identical in the test plant (II) and in the full-scale plant (I). Such parameters are bed temperature, total excess-air ratio, primary stoichiometry, fuel, and bed material. The fluidization velocity in the test plant should be in the same order of magnitude as in the full-scale plant. The fluid dynamic similarity (or lack thereof) will be discussed further below. It remains to determine the size of the test riser that is to represent a given full-scale combustion chamber. In some cases4,5 the time needed for relevant gas phase reactions has been chosen as a criterion to determine the height of the test riser. The width of the plant, on the other hand, cannot be scaled. The width has simply to be kept small. Otherwise the test plant would be converted into a pilot plant. Theory The formulation of the transport and reaction mechanisms is extremely complex, and an effort to follow the detailed processes would not lead to scaling but to (4) Kno¨big, T.; Werther, J.; Åmand, L.-E.; Leckner, B. Fuel 1998, 77, 1635-1642. (5) Sage, P. W.; Schofield, P. A.; Gaya´n, P. Proceedings of the 14th International Conference on FBC, F. D. Preto, S. Ed.; ASME: New York, 1997; pp 847-856.

(1)

Transport time is the time needed to transport the reactants for reaction. Instead of the four original Damko¨hler numbers, a vertical (Dav) and a horizontal (Dah) number will be used for the reaction processes in a fluidized bed combustor. In the vertical direction, the transport time is the time required for the vertical transport of the gases from the bottom, for instance, from the secondary air inlet to the height H above the bottom, or to the top of the combustor H ) Ho. In the case of fly char the corresponding height would be nHo, where n is the number of times the char particle of a given size passes the combustor. Considering the use of similar fuel, bed material, and temperature in the test plant and in the boiler to be modeled, the reaction times should be identical (if the mixing conditions are assumed to be similar) and only the transport times need to be considered. The average gas residence times

τ ) HI/uI ) HII/uII

(2)

yield a relationship between height and fluidization velocity u,

uII ) HIIuI /HI

(3)

which is of importance for the choice of the height of the plant HoII if the transport time is smaller than the reaction time (Dav < 1). The adjustment of the fluidization velocity may appear to be in conflict with the requirement of equal total excess air ratios. However, this is not the case as long as the fuel mass flow can be adapted. This latter approach was used4 in a comparison between the progress of reactions in a laboratory CFB and a pilot plant CFB boiler. It was also proposed4 to adjust the average solids concentration in the test plant to become similar to that in the boiler cjsII ) cjsI in order to have similar vertical distributions of solids in the two reactors. The measured total pressure drop of the riser can express this

∆p ) FsgHo cjs

(4)

∆pII ) ∆pIHoII/HoI

(5)

which gives

since the density of the particles is FsII ) FsI for the same bed material. The two important reactions in the horizontal direction, devolatilization and char combustion, take place while the fuel particles are dispersed horizontally from the fuel feed point (normally located at the wall of the combustion chamber) over the cross section of the furnace. If the time needed for this transport is small (6) Damko¨hler, G. Z.Elektrochem. 1936, 42, 846-862.

1288

Energy & Fuels, Vol. 14, No. 6, 2000

Leckner et al.

compared to the conversion time, that is, if Dah < 1, there will be an even distribution of fuel (volatiles and char) to meet the even distribution of oxygen from the air distributor. If Dah . 1, the fuel will not be evenly distributed, and there will be a deficit of oxygen on the feed side of the combustor’s cross-section and an excess of oxygen on the other side. Usually, the transport time is small in a narrow test unit but large in a wide commercial plant. This causes obvious problems of scaling, since, although the test plant can be tall to fulfill the height criterion, it can never be wide. Let us look more closely at the relationships for the horizontal scaling. Devolatilization time for a fuel particle of size d is commonly expressed as

τ ) adn

(6)

where a and n are empirically determined coefficients. Here, for simplicity a ) 106 s/m2 and n ) 2 will be used,7 although typical values8 could be a ) 1.6 × 106 s/m1.6 and n ) 1.6. The Nusselt square-law burning time for diffusioncontrolled char combustion can be written

τ)

Fcd2 4φMcShDCo

although it is recognized that it is uncertain for CFB bottom beds. The average dispersion distance x is used in Einstein’s classical expression for dispersion time10

τ ) x2/(2Dh)

(7)

The symbols and relevant values are as follows

Fc ) 1130 kg/m3, density of char (typical value) φ ) 1, reaction mechanism factor, the value 1 is chosen for a particle in a fluidized bed Mc ) 12 kg/kmol, molecular mass of carbon Sh ) 2, Sherwood number for a sphere at low Reynolds number -4

Figure 2. Horizontal SO2 and O2 profiles in a commercial boiler burning bituminous coal. The fuel is fed from left side. Adapted from Alliston and Wu.13

(8)

For a high volatile fuel the reaction time is governed by devolatilization, and the Damko¨hler number becomes (using eqs 6 and 8)

Dah )

( )[ ] x d

2

1 2aDh

(9)

or with the values given above

x (1000d )

2

Dah ) 50

(10)

Dah e 1 is valid for

2

D ) 2.25 × 10 m /s, molecular diffusion coefficient in the vicinity of the particle Co ) 0.53 × 10-3 kmol/m3, the oxygen concentration in the bed, assumed to be 5% There is some information in the literature on the dispersion of solids in small fluidized beds operating at low velocities. In contrast, the present interest is in large, high velocity fluidized beds. The only value of dispersion coefficient available for such situations is from a commercial size fluidized bed, operated under reasonably close conditions to the application concerned is, as far as known to the present authors, that of Bellgardt et al.,9 derived indirectly from combustion data obtained in a 5.5 × 3.7 m bed operated at 2.4 m/s. The fuel dispersion coefficient in this bed was Dh ) 0.01 m2/s. This value will be used in the present estimates, (7) Essenhigh, R. H.; Thring, M. W. Proceedings of the Conference on Science in the Use of Coal; Institute of Fuel: London, 1958; Paper 29. (8) Zhang, J. Q.; Becker, H. A.; Code, R. K. Proceedings of the 9th International Conference on FBC; Mustonen, J. P., Ed.; ASME: New York, 1987; pp 1203-1210. (9) Bellgardt, D.; Schoessler, M.; Werther, J. Powder Technology 1987, 53, 205-216.

(x/1000d) e 0.14 For a high-volatile fuel of a size of 1 mm the criterion is valid only for combustion chambers, having a diameter of less than 0.14 m, or for a 10 mm fuel, 1.4 m. Under such conditions it is uncertain to compare results from a 0.1m diameter research plant with a commercial plant, having a size on the order of 5 m, i.e., scale-up from the bench to the commercial scale is not possible. Maldistribution of fuel and air will be a characteristic feature of the large plant, but not necessarily of the small one. With char as a fuel, the reaction time is given by eq 7, and it follows

Dah )

( )[ x d

2

]

4McDCo F c Dh

(11)

With the values given above it holds

Dah ) 0.5

x (1000d )

2

(12)

(10) Gardiner, C. W. Handbook of Stochastic Methods, 2nd ed.; Springer: Berlin, 1997.

Fluidized Bed Combustion

Energy & Fuels, Vol. 14, No. 6, 2000 1289

Figure 3. Concentrations of O2, CO and NO on the center-lines of the CFB risers of CTH and TUHH.4

Dah e 1 is fulfilled if

(x/1000d) e 1.4 Obviously, x should be less than 1.4 m for a 1 mm char particle. This means that results from research plants should be transferable to the industrial scale. In the above examples x is the lateral mixing length. It was compared with a characteristic dimension of the combustion chamber, the (equivalent) diameter. However, the characteristic dimension can be chosen in different ways depending on the configuration studied: it could, for instance, be the distance between fuel feed points in a large combustor.

In application of fluid dynamic scaling criteria compromises have to be made: it is not possible to completely fulfill the criteria. The combustion scaling should be understood in the same way. The Damko¨hler criteria given can only be regarded as indications on feasible scaling. Furthermore, in the case of combustion the formal fluid dynamic scaling cannot be satisfied and has to be abandoned. The indicative nature of the criteria proposed is obvious. The fact that formal fluid dynamic scaling is neglected and that fluid dynamics instead are represented by identical average solids volume concentrations and order-of-magnitude similarity of fluidization velocities,

1290

Energy & Fuels, Vol. 14, No. 6, 2000

Leckner et al.

implies that the detailed fluid dynamics are not considered to have a decisive importance for the chemistry. This assumption may be reasonable in the upper parts of CFB riser, but the differences in fluid dynamic behavior in the bottom bed of a narrow test tube and a wide, commercial combustor, as mentioned above, may have a significant influence on the fuel conversion and pollutant formation. The consequences for the chemistry of the approximations regarding fluid dynamics are still not well-known. Results and Discussion The scaling procedure can be better conceived if illustrated by examples. Example 1: Desulfurization. Comprehensive tests have been carried out5 to investigate the feasibility of burning gasification residues in an atmospheric CFB combustor. The input fuel was char mixed with calcium sulfide produced by partial gasification of bituminous coal. The laboratory CFB combustor used was 8.7 m tall and had a diameter ranging from 0.15 m in the bottom part to 0.225 m in the top. Several tests were carried out within the ranges: bed temperature 900-950 °C, total excess-air ratio 1.06-1.4, and fluidization velocity 3 m/s, using several fuel and limestone brands with a size < 250 µm. The size of the char was around 1 mm, with a rather wide size distribution. A high combustion efficiency was obtained, and desulfurization was as high as 95 to 99% at a calcium to sulfur molar ratio of 2. The latter is remarkable, considering the high bed temperature and the common results from fluidized bed combustors that sulfur capture has a maximum at 850 °C and then decays at higher temperatures to become quite low at temperatures above 900 °C. Are then the laboratory tests representative? Sulfur capture is sensitive to the combustion condition in a furnace. Thermodynamics reveal that the conversion of CaO with SO2, as well as that of CaS, goes toward CaSO4 if sufficient oxygen is present. Under oxygen-deficient conditions, which could occur locally even at overall air excess, reverse reactions yielding SO2 may take place, especially in the presence of CO and H2, in combustors operating at atmospheric pressure. This may occur in the particle phase of a dense bed because of local oxygen depletion,11 or it may occur due to gross maldistribution of fuel and air. Both may give rise to the observed temperature dependence of sulfur capture.12 In the case of Sage and co-workers5 char was burned in a bed operated at the high velocity of 3 m/s, and the oxygen depletion in the dense bed should not be severe, because of the absence of volatiles and the considerable stirring of the bed by fluidization. Most important, however, the laboratory bed has a Dah < 1 and the mixing between fuel and air should have resulted in a prevailing excess of oxygen, even locally. This circumstance is most likely the explanation for the high sulfur capture efficiency in this case. Alliston and Wu13 conducted tests in a small laboratory reactor as a basis for design of boilers. Later, desulfurization results from the laboratory reactor were compared with those from the boilers, using the same fuels and limestones. They found that for low volatile fuels the agreement between desulfurization in the

Figure 4. Oxygen concentration on the center-line of the CTH riser (at 0.8 m from the walls) compared to the concentrations measured at 0.4 m from the walls where secondary air was injected.4

small and large scale plants was reasonably good, but for bituminous coals the test plant tended to give much better desulfurization than the large boilers. The behavior of the commercial boilers burning bituminous coal can be tentatively explained by the results of Alliston and Wu,13 Figure 2. This figure illustrates well the scaling problem related to the horizontal distribution of fuel: in the vicinity of the fuel feed point at the side of the combustion chamber oxygen is consumed by the volatiles. A reducing zone is created and SO2 is released from CaSO4 present in the bed (originating from sulfur previously captured in oxidizing regions) and produces high SO2 concentrations in this oxygenstarved region. Such an insufficient mixing between fuel and oxygen was not present in the test plant and consequently desulfurization was better there. The conclusion is again: scaling may be achieved in the vertical direction, but it is not possible to obtain a reasonable scaling for the width of a commercial plant, except for very low-volatile fuels. Example 2: Comparison between Plants. The criteria (3) and (5) were used in a comparison between a narrow laboratory combustor (TUHH; diameter 0.1 m, height 15.5 m) and a CFB boiler (CTH; cross-section 1.6 × 1.6 m, height 13.5 m), using three kinds of fuel, the same fuels in both units: coal, peat and wood.4 Some data are given in Table 1 where the impact of volatiles was simply weighted as Table 1. Fuel Characteristics and Horizontal Damko1 hler Numbers (Example 2) fuel

fraction of volatiles ξ

wood peat coal

0.78 0.70 0.39

fuel size, mma,b CTH TUHH 6 60 6

6 3 3

CTH 20 0.9 10

Daha TUHH 0.6 1.1 0.6

a These are sizes of the fuel fed to the bed. The size of the fuel in the bed is always smaller than that fed; this is especially the case of CTH peat. b The fuel had to be crushed before feeding into the smaller combustor.

(11) Dennis, J. S.; Hayhurst, A. N. Proceedings of the 20th Symposium (International) on Combustion; The Combustion Institute: Pittsburgh, 1985; pp 1347-1355. (12) Lyngfelt, A.; Leckner, B. Chem. Eng. Sci. 1989, 44, 207-213. (13) Alliston, M. G.; Wu, S. CFB Technology; Science Press: Beijing, 1996; pp 327-332.

Fluidized Bed Combustion

Energy & Fuels, Vol. 14, No. 6, 2000 1291

Figure 5. Geometries of the large sewage sludge incineration plant (4.5m diameter at grid level, 6.0 m diameter in freeboard) and of TUHH's small-scale combustor (0.10 m diameter at grid level and 0.15 m in freeboard).14

Dah ) ξ Dah,volatiles + (1 - ξ) Dah,char The table shows that Dah . 1 for the CTH boiler but not for the narrow laboratory riser, despite of the smaller fuel size. This means that there is a considerable lateral variation in the distribution of fuel and air in the CTH case, whereas the distribution is more even in the TUHH case. Obviously the two reactors are not scaled in the horizontal direction, but they are roughly scaled in the vertical direction. The measurements presented were made on the center-lines of the two risers, Figure 3. The overall emission values of the two plants agree reasonably well, but the local center-line concentrations can be quite different. It appears, judging from the oxygen concentration, that the oxygen is more rapidly consumed in the narrow TUHH riser than in the CTH riser. This is also in agreement with the higher CO concentrations in the bottom part of the CTH unit and indicates that combustion (mixing) is less efficient in this unit than in the smaller risersprobably an effect of the different fluid dynamics in addition to the transport aspect mentioned above. The rising oxygen concentration with height in the CTH unit differs from the distinct increase in oxygen concentration as a consequence of secondary air injection in the narrow riser (secondary air was used in a similar way in both units). The explanation is a gradual penetration of secondary air, which is not noticed in the small cross section of the narrow riser, but which plays an essential role in a larger riser, illustrated in a more detailed way in Figure 4. Although the height aspect of scaling was observed and the stoichiometries were the same, again the differences in cross-section area caused an additional difference between the two plants. In fact, it is a goal of design of large plants to eliminate this kind of differences: one would like the secondary air to penetrate perfectly and yield an even supply of oxygen at the level of injection.

Table 2. Comparison of the Emissions from Two Sludge Combustors14 (Example 3) TUHH InfraServ

NO [mg/m3]

CO [mg/m3]

NH3 [mg/m3]

90 122