Epstein N., Grace J.R. Spouted and Spout-Fluid Beds: Fundamentals and Applications

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The Wurster coater 247

Draft tube walls

Spray zone

Funneling

plate

Distributor plate

Nozzle

(a)

1.0-mm holes

2.5-mm holes

Deflector

Nozzle

Distributor plate

(b)

Figure 14.5. Schematic diagrams of (a) funneling plate used by Shelukar et al.

; and (b) tablet

deﬂectors used by Subramanian et al.

Chapter-speciﬁc nomenclature

N number of passes through spray zone

coat

total coating time, s

mass of coating added in ith pass through spray zone, kg

μ mean value

σ standard deviation

Subscripts

ct coating time

n number of passes

x per pass

248 Sarah Palmer, Andrew Ingram, and Jonathan Seville

References

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2. J. Hogan. Coating of tablets and multiparticulates. In Pharmaceutics: The Science of Dosage

Form Design, ed. M. E. Aulton (London: Churchill Livingstone, 2002), pp. 441–448.

3. M. Tzika, S. Alexandridou, and C. Kiparissides. Evaluation of the morphological and release

characteristics of coated fertilizer granules produced in a Wurster ﬂuidized bed. Powder

Technol., 132 (2003), 16–24.

4. G. Subramanian, R. Turton, S. Shelukar, and L. Flemmer. Effect of tablet deﬂectors in the

draft tube of ﬂuidized/spouted bed coaters. Ind. Eng. Chem. Res., 42 (2003), 2470–2478.

5. E. Teunou and D. Poncelet. Batch and continuous ﬂuid bed coating – review and state of the

art. J. Food Eng., 53 (2002), 325–340.

6. D. Jones. Air suspension coating. In Encyclopedia of Pharmaceutical Technology,ed.J.

Swarbrick and J. Boylan (New York: Dekker, 1988), pp. 189–216.

7. D. E. Wurster. Means for applying coatings to tablets or like. J. Amer. Pharmaceutical Assoc.,

48 (1950), 451.

8. K. Jono, H. Ichikawa, M. Miyamoto, and Y. Fukumori. A review of particulate design for

pharmaceutical powders and their production by spouted bed coating. Powder Technol., 113

(2000), 269–277.

9. X. X. Cheng and R. Turton. The prediction of variability occurring in ﬂuidized bed coating

equipment. I. The measurement of particle circulation rates in a bottom-spray ﬂuidized bed

coater. Pharmac. Dev. & Technol., 5 (2000), 311–322.

10. D. Geldart. Types of gas ﬂuidization. Powder Technol., 7 (1973), 285–292.

11. T. Crites and R. Turton. Mathematical model for the prediction of cycle-time distributions for

the Wurster column-coating process. Ind. Eng. Chem. Res., 44 (2005), 5397–5402.

12. F. N. Christensen and P. Bertelsen. Qualitative description of the Wurster based ﬂuid bed

coating process. Drug Develop. & Ind. Pharm., 23 (1997), 451–463.

13. R. Turton and X. X. Cheng. The scale-up of spray coating processes for granular solids and

tablets. Powder Technol., 150 (2005), 78–85.

14. S. E. Palmer. Understanding and optimisation of pharmaceutical ﬁlm coating using the

Wurster coater. Ph.D. thesis, Univ. of Birmingham, UK (2007).

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recycle systems, AIChE J. 25 (1979), 873–882.

16. U. Mann. Analysis of spouted-bed coating and granulation. 1. Batch operation. Ind. Eng.

Chem. Proc. Des.Dev., 22 (1983), 288–292.

17. X. X. Cheng and R. Turton. The prediction of variability occurring in ﬂuidized bed coating

equipment. II. The role of nonuniform particle coverage as particles pass through the spray

zone. Pharm. Dev. & Technol., 5 (2000), 323–332.

18. S. Shelukar, J. Ho, J. Zega, E. Roland, N. Yeh, D. Quiram, A. Nole, A. Katdare, and S.

Reynolds. Identiﬁcation and characterization of factors controlling tablet coating uniformity

in a Wurster coating process. Powder Technol., 110 (2000), 29–36.

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Chem. Eng., 53 (1975), 579–581.

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on the ﬁlm thickness of beads coated in ﬂuidized bed equipment. Int. J. Pharmaceutics, 65

(1990), 69–76.

The Wurster coater 249

21. M. Choi. Hydrodynamics of commercial-scale Wurster coaters. Presented at annual AIChE

meeting (1998).

22. S. Fitzpatrick, Y. Ding, C. Seiler, C. Lovegrove, S. Booth, R. Forster, D. J. Parker, and

J. P. K. Seville. Positron emission particle tracking studies of a wurster process for coating

applications. Pharmaceutical Technol., 27 (2003), 70–78.

23. J. P. K. Seville, U. T

un, and R. Clift. Processing of Particulate Solids (Glasgow: Chapman

and Hall/Blackie, 1997).

24. R. Turton. Challenges in the modelling and prediction of coating of pharmaceutical dosage

forms. Powder Technol., 181 (2008), 186–194.

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experimental and modelling studies. Powder Technol., 74 (1993), 259–270.

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uniformity in a Wurster ﬂuidized bed coating process. Powder Technol., 166 (2006), 81–90.

28. C. Seiler, P. J. Fryer, and J. P. K. Seville. Statistical modelling of the spouted bed coating

process using positron emission particle tracking (PEPT) data. Can. J. Chem. Eng., 86 (2008),

571–581.

29. P. W. S. Heng, L. W. Chan, and E. S. K. Tang. Use of swirling airﬂow to enhance coating

performance of bottom spray ﬂuid bed coaters. Int. J. Pharmaceutics, 327 (2006), 26–35.

15 Gasiﬁcation, pyrolysis, and

combustion

A. Paul Watkinson and Antonio C. L. Lisboa

A common set of reactions, given in Table 15.1, occurs when carbonaceous solids

undergo thermal processing. Whether the aim is pyrolysis, gasiﬁcation, or combustion,

each of these reactions occurs in some parts of the reactor because of gas–solids con-

tacting. In this chapter, we consider, in order, gasiﬁcation as an endothermic process to

generate H

and CO mixtures for fuel or synthesis gas, pyrolysis as a process to generate

useful tars (or liquids), and combustion as a process to produce heat.

15.1 Gasiﬁcation background

Most commercial gasiﬁers use coal as feed, and may be classiﬁed by the type of solids–

gas contacting (moving, entrained, ﬂuidized, or spouted bed), by the state of the ash (dry,

agglomerated, or molten), and by the oxidant (air, air–steam, or oxygen–steam). A low-

caloriﬁc-value gas results from air–steam gasiﬁcation, and a medium-caloriﬁc-value gas

from using steam or steam–oxygen mixtures. The carbonaceous feedstock may be fed

as a dry solid, a sludge, or a slurry.

Performance measures can be identiﬁed for comparing gasiﬁers of different designs

and operating conditions. For production of fuel gases, the heating value of the pro-

duced gas is important and is usually reported on a dry gas basis. For synthesis gas

or pure hydrogen production, the molar ratio of H

/CO leaving the gasiﬁer is criti-

cal. Low tar yields are usually beneﬁcial, unless a raw fuel gas is desired. For sizing

scaled-up gasiﬁcation processes, throughput of solids feed per unit cross-section of

reactor (kg/m

s) is of major importance. A high gas velocity in the gasiﬁer is beneﬁ-

cial, as it usually corresponds to a high solids throughput. A high carbon conversion

into reducing gases – excluding CO

– by the reactions of Table 15.1 is of great-

est importance in single-stage gasiﬁcation. When gasiﬁcation is followed by a sec-

ond stage of char combustion (for example, to produce electrical power), high carbon

conversion in the gasiﬁcation reactor is of less importance. In such cases, mild par-

tial gasiﬁcation or devolatilization of the feed with low conversion of carbon may be

desired.

Spouted and Spout-Fluid Beds: Fundamentals and Applications, eds. Norman Epstein and John R. Grace.

Published by Cambridge University Press.



Cambridge University Press 2011.

Gasiﬁcation, pyrolysis, and combustion 251

Table 15.1. Subset of simpliﬁed reactions in gasiﬁcation or combustion systems.

∗

Pyrolysis Carbonaceous solid → char +tar +gases

,CO,CO

O, CH

)

Neutral

Gasiﬁcation C (char) +H

O → CO +H

C(char)+CO

→ 2CO

C(char)+2H

→ CH

Endothermic

Combustion C (char) +O

→ CO

C(char)+

→ CO

Exothermic

Gas phase CO +H

O → CO

O → CO +3H

CO +

→ CO

Exothermic

∗

Ignores tar cracking, gas phase combustion, reforming of C

and tars,

reactions of N, S, minerals in feed solid.

15.2 Spouted bed gasiﬁcation of coal and coke

Spouted beds have been used to gasify coal, coke, carbonaceous sludges, biomass, and

municipal solid waste. In most cases the feed material has been coal, and the spouted bed

gasiﬁer has been operated at atmospheric pressure in dry ash mode with an air–steam

mixture as the gasifying agent. Some authors have reported studies at elevated pressures,

with oxygen–steam mixtures, or in the ash-agglomerating mode. For small-scale

(<0.25-m diameter) units, electrical furnaces can provide heat for the gasiﬁcation

through the reactor walls; otherwise, operation can be autothermal, with the heat gener-

ated in situ by combustion reactions driving the endothermic char–steam or char–CO

reactions. A variety of feeding arrangements have been used. Coal particles, typically a

few millimeters in diameter, can be added on top of the spouted bed, through the side

of the reactor, at the apex (bottom) with the entering oxidant gas ﬂow, or in other ways,

as shown in Figure 15.1. In attempts to develop spouted bed gasiﬁcation technology,

various conﬁgurations such as normal spouting, spout ﬂuidization, and spouting with

draft tubes have been explored.

From the early 1980s, publications appeared on spouted bed gasiﬁcation of coal mainly

from Canada, the United Kingdom, Japan, and the United States. Recent work has come

from China, where most of t he world’s coal gasiﬁcation systems are found.

Table 15.2

summarizes equipment and operating conditions, together with some key results.

After reporting initial results from a 0.15-m diameter reactor,

Foong et al.

and

Watkinson et al.

4,5

carried out most of their subsequent work in a 0.31-m diameter,

autothermal atmospheric pressure spouted bed gasiﬁer, with coal feed rates up to 60 kg/h.

Although top feeding was t ested, in most work coal of size 0.7 to 3.5 mm was conveyed

into the bottom of the reactor (Figure 15.1a) with the spouting air–steam mixture. Ash

was removed from the produced gas and recovered by cyclone. The iner t bed was

composed of crushed gravel of 1 to 3.4 mm diameter. The bed reached a steady-state

carbon content and temperature for autothermal operation, depending on coal feed rate

and reactivity. The spouted bed gasiﬁer was shown to be operable with coals of all ranks,

including caking coal, as well as with delayed coke.

With caking coal, top-feeding

252 A. Paul Watkinson and Antonio C. L. Lisboa

GAS

SOLIDS

FEED

SPOUTING GAS

CHAR/ASH

TOP

SIDE

APEX

(a)

GAS

CHAR

+ H

CHAR

AIR

FUEL

COAL

COMBUSTOR

SLAG

(b)

GAS

ASH/CHAR

STEAM

FEED

DEFLECTOR

ASH/CHAR

(c)

FEED

ASH, SORBENT

SPOUT/FLUID

GAS

SORBENT

(d)

Figure 15.1. Solids and gas feeding arrangements for different gasiﬁer conﬁgurations: (a) spouted

bed, adapted from Watkinson et al.

3–7

; (b) agglomerating ash gasiﬁer, adapted from Kikuchi

et al.

; (c) jet-spouted bed, adapted from Tsuji et al.

9–11

; (d) spout-ﬂuid bed, adapted from Arnold

et al.

should be avoided. Throughputs and carbon conversion were both higher for more

reactive, lower-rank, noncaking coals. Outlet gas quality and yield were reported as

functions of such properties as coal feed rate, particle size, bed temperature, and oxidant-

to-coal ratio. Results were obtained for single pass versus char recycle,

with oxygen–

steam versus air–steam blowing in agglomerating versus dry-ash operation, with addition

of catalyst,

and with addition of limestone and dolomite for sulfur capture.

In cases

in which an inert bed was not used, the resulting char bed was less stable, with the bed

depth undergoing signiﬁcant changes over time. In-bed temperature and composition

proﬁles were measured.

Figure 15.2 shows that oxygen is limited to the lower regions of

the spout; higher up, the spout is under reducing conditions. Oxygen is virtually absent

from the annulus, which is largely a reducing zone, with highest concentrations of CO

and H

at the top of the bed. The temperatures reﬂect the gas analyses, with the oxidation

in the spout resulting in temperatures at the top that exceed those in the annulus. Such

data provided inputs for modeling studies.

Kikuchi et al.

developed an agglomerating ash gasiﬁer system using the conﬁguration

shown in Figure 15.1b, with an upper spouted gasiﬁer of D = 0.4 m, operated with

oxygen–steam mixtures injected both below the conical section and two-thirds of the

way up the cone height. Up to 100 kg/h coal was fed through the side of the reactor into

Gasiﬁcation, pyrolysis, and combustion 253

Table 15.2. Summary of spouted bed gasiﬁcation studies by reactor size and throughput.

∗

Feed

Reactor

type D (m)

Feed rate

(kg/h) Oxidant

max

(MPa) T (

◦

Max. solids

ﬂux (kg/m

-s)

Reference

No.

coal S 0.028 0.16 O

-N

1.3 850−900 0.07 21

coal S 0.076 − N

+St 2 900 − 23

coal SF 0.08 5−12 A +St 0.5 870−1030 0.66 17

biomass S, F 0.09 4.2 A 0.1 607−842 0.18 28

coal, an S 0.10 5 St, O, A 1.8 550−865 0.18 14−15

c.coal S 0.13 10−16 A 0.1 980−1185 0.33 27

coal S 0.15 2.5−12 A +St 0.1 750−930 0.20 2

coal S, J 0.21 4−8A+St 0.1 727−1027 0.06 9

coal J 0.20 7−10 O +St 0.1 1007−1120 0.09 11

biomass SF 0.26 80 A 0.1 700−765 0.42 24

coal S, F 0.305 60 A +St 0.1 805−985 0.23 3, 4

coal S 0.305 28−42 O +St 0.1 710−1120 0.16 5

coke S 0.305 16−41 O +St 0.1 900−970 0.16 6

coal S 0.4 50−100 O +St 0.1 1050−1170 0.17 8

coal SF 0.4 360−490 A +St 1.8 970−1020 0.90 13

coal SF 0.45 317−330 A +St 0.5 940−980 0.60 18

msw SF 0.45 ×0.45 ∼70 A 0.1 662−777 0.10 26

msw S 1.0 700 O +St 0.1 540−870 0.25 25

coal SF 1.2 500 A

+St 0.1 950−1020 0.12 12, 13

coal S 1.0 2490 A +St 1.2 980 0.90 19

∗

Feed: an = anthracite, c.coal = charcoal, msw = municipal solid waste

Reactor type: S = spouted, SF = spout-ﬂuid, F = ﬂuidized, J = jet spouted

Oxidant: A = air, St = steam, O = oxygen

H (m)

0.0 0.2 0.4 0.6

Gas temperature (°C)

650

700

750

800

850

H (m)

0.0 0.2 0.4 0.6

Gas composition (vol.%, dry)

H (m)

0.0 0.2 0.4 0.6

Balance N

Figure 15.2. Axial composition and temperature proﬁles in the near-wall annulus A, mid-annulus

B, and spout C, adapted from Haji-Sulaiman et al.

254 A. Paul Watkinson and Antonio C. L. Lisboa

Carbon conversion (-)

Circles: F

= 6.48 kg/h

Tr ia ngles: F

= 10.4 kg/h

0.8

0.9

= 6.48 kg/h

= 10.4 kg/h

Oxy

en/coal ratio (k

)

0.55 0.60 0.65 0.70 0.75 0.80 0.85

Maximum bed temperature (°C)

800

900

1000

1100

1200

Lower stage

Upper stage

Lower stage

Upper stage

= 6.48 kg/h [

= 10.4 kg/h [

Gas composition (vol.%)

Figure 15.3. Effects of oxygen/coal mass ratio on outlet gas composition, carbon conversion, and

gasiﬁer bed temperatures in a jet-spouted bed, adapted from Tsuji and Uemaki

= coal feed

rate.

a relatively shallow bed. Temperatures were 1050

◦

C to 1170

◦

C, at which 2- to 10-mm-

diameter ash agglomerates of low (<4%) carbon content formed, and fell through the

throat into a pool of slag, into which cyclone char was injected. The effects of key

process variables were presented. Uemaki and Tsuji

9–11

describe a dry-ash, jet-spouted

bed (Figure 15.1c) oxygen–steam-fed gasiﬁer system of D = 0.2 m with a coal feed

rate of ∼10 kg/h. In this case, a sharper cone angle (θ = 40

◦

) and a shallow bed system

(H/D ∼ 2) were used at velocities of 0.8 to 1.8 m/s. This resulted in a high bed voidage

(ε ∼0.9), which permitted operation up to 1150

◦

C without ash agglomeration. Relatively

uniform axial temperature proﬁles were found. Effects of key process variables on

gas quality, gas yield, and carbon conversion were explored. Figure 15.3 shows the

interplay of oxygen–coal ratio, temperature, carbon conversion, and gas quality typical

of gasiﬁcation processes. As the O/C ratio is raised with a ﬁxed ﬂow of coal, temperatures

and carbon conversion values rise, and the H

/CO ratio in the produced gas declines.

Because the heats of combustion per mole of hydrogen and carbon monoxide are very

similar, the gas heating value increases only weakly with increasing O/C ratio.

British Coal developed an atmospheric pressure spout-ﬂuid bed gasiﬁer for low-

caloriﬁc-value gas

12,13

in which most of the air–steam mixture is introduced at the

apex of the conical section, and the remainder along the cone (Figure 15.1d). For a bed

Gasiﬁcation, pyrolysis, and combustion 255

deeper than H

, the upper portion behaves as a slugging ﬂuidized bed. Units have been

operated at up to 500 kg/h with different coals, feed sizes up to 25 mm, different bed

materials and bed depths, with air enrichment by oxygen, and with limestone injection.

The beneﬁts of operating with a dense, inert bed of solids (bauxite), rather than a char

bed, were demonstrated. With some coals, unwanted agglomeration occurred at bed

temperatures above 1000

◦

C in silica sand beds. Carbon conversion up to 94 percent

was possible with char recycle. Gas heating values were about 3.9 MJ/m

. Eighty-ﬁve

percent desulfurization of the gas was achieved. Scaleup issues for the atmospheric

pressure unit to a 0.5-m diameter gasiﬁer have also been discussed.

A pressurized

spout-ﬂuid gasiﬁer unit, capable of throughputs of 490 kg/h, at pressures to 1.8 MPa with

a 0.32-m-diameter × 10-m tall pressure vessel, had been fabricated and commissioned

by 1992.

Pressurized spouted bed gasiﬁcation in a 0.1-m-diameter, 2–5 kg-coal/h gasiﬁer blown

with steam or steam–oxygen mixtures, and able to operate at pressures to 1.8 MPa, was

described by Sue-A-Quan et al.

14,15

When ﬁring with steam alone, pressure effects on

gas composition were minimal, and carbon conversion was shown to increase linearly

with pressure up to ∼1.5 MPa before leveling out. Oxygen addition was necessary

when using anthracite because of its low reactivity. Catalysts were shown to be effec-

tive in increasing carbon conversion at a given temperature. A gasiﬁer based on a

draft tube spouted bed (Chapter 7) was described in a series of papers by Hatate and

co-workers.

Recent work from China by Xiao et al.

featured an 0.08-m diameter spout-ﬂuid

unit operated at a pressure of 0.5 MPa with a high-temperature (700

◦

C) inlet air–steam

mixture to gasify up to 12 kg/h of high-ash coal at 870–1030

◦

C. Pressure and temperature

effects are delineated, with carbon conversion in the absence of char recycle reaching

about 86 percent at the highest temperature. The dense bed was essentially isothermal.

Spout temperatures were 20–30

◦

C higher than in the annulus, in rough agreement with

previous work.

In subsequent research, Xiao et al.

developed a spout-ﬂuid bed pilot

plant that is said to be insensitive to the high ash levels in their low rank coals. This

0.45-m diameter × 10.5-m-high air–steam blown gasiﬁer was designed for operation

at 0.5 MPa and 950–980

◦

C. Bed inert material is 1.85-mm-diameter boiler ash, which

provides a deep bed (3.6 m depth), in a manner similar to the British Coal process.

Coal fed at 330 kg/h is gasiﬁed using preheated air–steam mixtures at a superﬁcial

velocity of 1.3 m/s. Carbon conversion efﬁciencies up to 69.4 percent are reported, with

produced gas heating values of 4.5 MJ/m

. In a combustor unit, bottom ash containing

up to 39 percent carbon is to be burned with 52 percent carbon content ﬂy ash from the

cyclones.

In the United States, the Power Systems Development Facility (PSDF) circulating

ﬂuid bed system under development in the early 1990s included a pressurized spouted

bed “carbonizer” or mild gasiﬁer.

Here a 1-m-diameter × 18-m-high spouted bed

treated 2490 kg/h of <3.2 mm coal mixed with 475 kg/h sorbent at a temperature of

980

◦

C and a pressure of 1.3 MPa, to produce a fuel gas of heating value 4.9 MJ/m

and

a char product for subsequent combustion.

256 A. Paul Watkinson and Antonio C. L. Lisboa

Sulfur capture from troublesome gases such as H

S and COS is mentioned above.

Ammonia formation during coal gasiﬁcation has been investigated

20,21

in a very-small-

diameter (D = 28 mm) spouted bed able to operate up to high (20 bar) pressures at a

coal feed rate of about 0.2 kg/h. NH

levels of up to 2750 ppmv were reported. The

same system was used to gasify dried sewage sludge of 36 percent ash

at 770

◦

Cto

790

◦

C and pressures of 0.2 to 0.4 MPa with air. The gas produced had heating values

of 1.6 to 5.1 MJ/m

, and, as well as considerable H

S, contained up to 8000 ppm NH

arising from the 4.5 percent by weight N in the feed. High temperatures led to reduced

ammonia concentrations.

Effects of coal particle size on both pyrolysis and steam gasiﬁcation at 900

◦

C in deep

beds of small (76 mm) diameter have been investigated.

Coal particle size effects are

also mentioned in other work.

3,12

15.3 Spouted bed gasiﬁcation of biomass and wastes

The extensive increase in gasiﬁcation research over the past decade or so, focused mostly

on biomass f eedstocks, has yielded relatively little recent work on spouted beds. Early

spout-ﬂuid bed gasiﬁcation of biomass wastes was reported by Zak and Nutcher.

1993, the s pouted bed reactor was described as an “emerging technology” for recycling

waste,

with solid wastes gasiﬁed at temperatures of 538–870

◦

C by highly super-

heated steam to produce a syngas product and an inert granular solid. Carbonaceous

feedstocks with heating values of 7 to 28 MJ/kg, including coal-tar-contaminated soils,

petroleum reﬁnery wastes, and municipal solid wastes, could be treated. A unit of

approximately 1 m diameter, with bed material of 1.6 to 6.4 mm silica sand, treated up to

680 kg wastes/h and produced a gas of 12 to 18 MJ/m

and an ash containing 1 percent to

2 percent carbon. Auxiliary fuel and oxygen were fed at the bottom; char was oxidized

in the cyclone to produce a molten inert ash. Gasiﬁcation of simulated municipal solid

waste in a spout-ﬂuid bed of square cross-section was described by Thamavithya

and Dutta

; feed rates and temperatures were not given. Salam and Bhattacharya

reported on spouted bed gasiﬁcation of wood charcoal with both a normal circular

gas inlet oriﬁce and a novel circular slit oriﬁce. Hoque and Bhattacharya

compared

ﬂuidized and spouted bed gasiﬁcation of coconut shells.

15.4 Throughput and scale of spouted bed gasiﬁcation

The scale to which spouted bed gasiﬁcation has been demonstrated is of key interest

in assessing the potential for industrial implementation. The previous survey covered

results from spouted bed units processing from less than 5 kg coal/h to 2490 kg/h,

a scale factor of more than 500 in terms of feed rate. Table 15.2 indicates that the

increased throughput has been achieved by expanding reactor diameters from 0.15 m to

1.2 m, a scaleup in cross-sectional area of about 60. By operating at 1.0 MPa rather than

atmospheric pressure, a further capacity increase of roughly a factor of ten occurs. It is