K3. Factors governing the behaviour of the adiabatic water-gas shift reactor

K3. Factors governing the behaviour of the adiabatic water-gas shift

reactor

P. MARS Central Laboratory, Staatsmijnen in Limburg, Geleen, Netherlands Ahstnrct-The temperature protik in an adiabatic water-gas shift reactor was calculated on the basis of laboratory measurements with the aid of a computer. The laboratory data used comprise: the intrinsic activity of the catalyst, the effect of sintering, the influence of diffusion in the catalyst pores and the poisoning etfect of hydrogen sulphide. The influence of mass transfer to the surface of the pellets can be neglected. The result agrees on the whole with that found in practice, ieaisnecessary. although further investigation of some dii The infhtence of the activity of the catalyst and of the particle size on the temperature profile and on the conversion was calculated in the same way. B-Auf Rasis von Laboratoriummessungen wurde unter Verwendung einer Rechenmaschine das Tempcraturprotil eines adiibatischen CO-Konvertors errechnet. Die hienu bent&ten Laboratoriumdaten waren: die Eigen-aktivitit des Katalysators, der Sintereffekt, der Eintluss von Diffusion in den Katalysatorporen und der Hemmungseffekt des Schwefelwasserstoffs. Der von der Stofft3ertragung ausgehende Einfiuss auf die Reaktionsgeschwindigkeit ist geri&%gig. Das !3gebnis der Rerechnung stimmt im grossen und ganzen mit dem in der Praxis gefundenen tiberein. obgleich es notwendig ist, einige Abweichungen zu untersuchen. In gleicherWeise wurde der Einfluss der Aktivitit des Katalysators und der Teikhengrosse auf das Temperaturprofil und auf die Konversion errechnet. R&ums6-A la base de mesures effect&a B 1’6chelle de laboratoite on a calcule la marche de la temp&atute B l’aide d’une machine a cakuler. A cet effet on a utili& les don&s de laboratoire suivantes: Pactiviti intrin&que du catalyseur, l’effet de fiittage du catalyseur, l’influence de la diffusion dans ks pores du catalyseur et l’effet d’empoisonnement par Pa&k sulfhydrique. L’iniluence du transfert de masse a la surface des globules eat &gligeable. Lea rrcsultatsdu calcul correspondent grandement a ceux trouves dans la pratique bien qu’il soit neccssaite d’examiner quelques &arts. De la m&me facon on a cakule l’influence de l’activite du catalyseur et de la grosseur de grain sur la marche de la temperature et sur la conversion.

INTRODU~TTION OUR

knowledge of the chemical and physical factors which govern the behaviour of technical reactors is not yet quite sufficient to design these exclusively on the basis of laboratory experiments. The greater part of the present large installations are designed semi-empirically, use being made of calculations of some details. To increase our knowledge it is useful to make calculations on actual technical reactors and to compare the results of these with laboratory experiments. In the paper presented here the practical data drawn from an adiabatic water-gas shift reactor filled with pellets of a normal iron oxide-chromium

oxide catalyst will be compared with the results of laboratory experiments. In the lirst part of the paper we shall discuss the determination of the activity of the catalyst in the laboratory. In the second part the data obtained will be compared with the activity data calculated from the temperature protie in the technical reactor. In the last section we shall report calculations of the behaviour of the reactor as a function of catalytic activity and reaction conditions. PART I.

LABORATORY EXPERIMENTS

Experimental procedue I*’ A diagram of the apparatus used is given in Fig. 1. A mixture of CO, COz and He is passed

375

P. MARS

FIG. 1. Diagram of laboratory apparatus for the determination for the water-gas shift reaction.

through a water vapour saturator, which is kept at a constant temperature of about 90 “C. A small amount of carbon-monoxide containing 1 per cent H2S is added to the gas stream, which is fed into the reactor. The reactor consists of a U-tube (dia. 8 mm), the lower part of which is filled with quartz grit (05 mm), which serves as a preheater. The catalyst (about 1.4 g, dia. 0.5-O-7 mm) is diluted with 5-12 ml of the same quartz grit. The reactor is surrounded by a fluid bed of Carborundum powder (200~) acting as a heating bath. In this way a very uniform temperature is obtained over entire length of the reactor. Heating is performed electrically on the outside of the fluid bed and is regulated with a platinum resistance thermometer dipping into the bed and connected to a normal electronic relay system. The temperature is measured with a thermocouple. The gases which leave the reactor are cooled and dried. After absorbing the CO2 on soda lime the CO/H2 ratio is measured with a catharometer, taking the feed gas (after elimination of COs) as reference gas. Space velocities were in between 10,000 and The degrees of conversion never 45,000 h-l. exceeded 20 per cent.

of the activity of catalysts

The reduction of the catalyst sample to Fe304 is performed with the normal gas mixture at 400 “C for 30 min. After the reduction period the testing temperature is adjusted. In most cases not more than 30 min at test conditions were necessary to reach a constant conversion. I.2

Transport restrictions

In order to measure the real chemical activity of the catalyst it is necessary to take into account the influence of the following physical factors on the experimental reaction rate : (a) Mass transfer to and from the geometrical particle surface. (b) Heat transfer from the catalyst to the gas phase. (c) Diffusion restrictions in the catalyst pores. The external mass transfer can be calculated with the aid of well-known relations between dimensionless numbers [l]. These calculations and also the more generalized considerations of WICKE [2] show that under our experimental conditions the rate of mass transport has no influence on the experimental reaction rates. The exothermic character of the reaction may increase the temperature of the catalyst particles

376

Factors

governingthe behaviour of the adiabaticwater-gasshift reactor

coefficient in the tablet Den was measured by the method of HOOGSCHAGEN [4]. Unlike this author [4, 51 we take the diffusion in the pores to be of the Knudsen-type. This is based on measurements of the diffusion rates of various gases through catalyst tablets by VAN HEERDENand NOBEL [6]. This is in agreement with the average pore diameter, which amounts to 300 A, and with the fact that the volume of the pores with a diameter greater than 1000 A is only 10 per cent of the total pore volume [7]. The chemical reaction rate is defined by the firstorder reaction constant k’. The (apparent) Thiele modulus

to above the temperature of the gas. It is well known that the temperature depression of a wet ball-thermometer is equal to the temperature decrease of the gas in the case of adiabatic vaporization of water up to saturation. An analogous phenomenon is observed in the case where the reaction rate on a catalyst particle is fully determined by mass transport; in that case the temperature difference between the gas phase and the catalyst particle is equal to the temperature increase caused by adiabatic equilibration. If mass transfer determines the reaction rate only partially, one may state as a rough approximation:* (1)

m’ =

$$k’/Diff)lf’

in which can thus be calculated, and from m’ we get the effectiveness factor Eff (= the ratio of actual reaction rate to the reaction rate not restricted by pore diffusion). Obviously Eff increases as Deft increases and as +,, or k’ decreases. The result of our calculations showed that for our experimental conditions Eff > 0.95 for measurements up to a temperature of 380 “C, and that consequently the chemical activity is determined directly.

= the temperature difference between particle and gas = the temperature increase of the gas caused AK* by adiabatic equilibration k eff =the reaction rate observed = mass transfer coefficient k, AT

Since ktir/km g 0.01 and ATM ‘Y 250 “C, it follows that AT g 2.5 “C, which can be neglected. Further calculations on the heat conductivity of granular beds and heat transport to the reactor wall showed that radial temperature gradients in the reactor tube are also negligible. The retarding effect of the diffusion of the reaction components in the catalyst pores can under certain conditions be very considerable. The influence of the diffusion of CO is the first to become perceptible, because pco has a great influence on the reaction rate. However, COs has a small retarding effect (see below), but its concentration is already appreciable. Ha and HsO have practically no influence on the reaction rate [3]. The influence of internal diffusion can be calculated by means of the formulae derived by THWLE and WAGNER,which we used in the form given by HOOGSCHAGFJN [4]. The effective diffusion

I.3 The kinetics of the reaction In or&r to be able to express the catalytic activity of a given sample in a reaction rate constant, some information about the kinetics of the reaction is necessary. In the absence of such information the catalytic activity can only be expressed in terms of the space velocity at a predetermined constant degree of conversion. The kinetics of the reaction under consideration is not completely elucidated. Some authors assume a fist-order reaction in CO [8,9], but other opinions are also found. STELLINGand VON KRUSENSTIERNA [3] showed that the reaction rate is proportional to PCO, but is lowered by COs. This has been confIrmed by our own experiments. It has been found that the first-order rate constant k, defined by

*This relation cau be deduced from the similarity of the k=-V,ln(l-q/rtcp) mechanism of heat and mass transfer, assuming that the Praudtl number is equal to the Schmidt number (HOWEN was a continuously decreasing function and WATSON, p. 982[ID. 377

(2) of

pcor.

P. Maas

Since all our experiments reported here have been performed with gas mixtures resembling those used in industrial practice, the amount of CO2 is always large and does not change appreciably during the reaction, so that the first-order constant k can be used directly as a measure of the activity of the catalyst. 1.4 Results The results of our experiments are given in Table 1 and can be described by the formula: k = k, exp(- E/RT), with E = 32,000 Cal/mole. Table 1. Results of the laboratory measurements of the catalytic activity of the iron oxide-chromium oxide catalyst Gas composition (without Ns and HaO) about 60 per cent CO, 10 per cent COS and 30 per cent He + a standard amount of sulphur

I

I

Temperature (“c)

Activity 01-3

360 335 315 380 360 335

7900 1780 910 10900 6440 2240

PART II.

THE DE~ERMNATION OF THB CATALYTIC Acrrvrr~ FROMTECHNICALDATA

II. 1 The reactors The conversion of water-gas is carried out in two consecutive adiabatic reactors. The feed gas is mixed with l-2 volumes of steam and passes through the reactors from top to bottom. In the first reactor the main part of the CO reacts and the gas approximately reaches equilibrium at the increased outlet temperature. The gas mixture is then cooled and in the second reactor the gas mixture approaches the more favourable equilibrium at the lower temperature. In the jrst reactor vertical tubes are present in which a thermocouple can be inserted, thus making it possible to measure the vertical temperature profile at different places in the be 1. In Figs. 3 and 4 two of these profiles are given.

Space velocity (h-9

r-------

7.1

46OLm 23800 9000

450 I

I

Apart from these measurements, the inlluence of varying amounts of HsS on the catalytic activity was also determined. No COS was added, because it appeared that under normal reaction conditions this component reacts very rapidly with Hz0 to form H2S. Fig. 2 shows the relative change of the catalytic loo

3,&----

Bona

FIG.3. Temperatureprofile in first mctor (at the beginning of a working period). --measured -.-.-

Sulphw concentmtian Fro. 2.

in the gas,

wbitmy

unik

The catalytic activity of the iron oxide catalyst as a function of the HsS concentration. 0 400 “C x 345 “C

activity as a function of an increasing HsS content.

>

calculated

It is obvious that under strictly adiabatic conditions these experimental temperature profiles provide sufficient information to calculate the catalytic activity, if all the factors influencing it are known. In the second reactor only the temperature over and underneath the catalyst bed is measured. Since 378

Factors governing the behaviour of the adiabatic water-gas shift reactor

(b) The temperature of the catalyst particles may be higher than the gas temperature, because of the exothermic chemical reaction. This effect was estimated along the same lines as adhered to for the laboratory experiments. For the conditions in the technical reactor the results indicated in Table 2 were obtained. Table 2. Injpuence of Mass and Heat Transfer

A

WJ

ken/km

400 450 500

0.010 0.025 0.053,

1.4 2.2 2.1

uollom

FIG. 4. Temperature profile in first reactor (after a working period of several years). --measured calculated

*This result is contradictory to the conclusionsof Boa-rom 1121, who says that external transport is exclusively ratb determining for the water-gas shift reaction at 500 “C. In our opinion, however, the insufficient temperature control in his experiments accounts for this contradiction.

these temperatures do not differ by more than 38 “C, it is possible to determine the first-order reaction rate constant at the average temperature directly from the decrease of the CO concentration in the gas. II.2

Temperature (“Cl

Reliability of the temperature profiles measured

Different factors may lead to errors in the temperature measurement in the technical reactor: (a) The heat conductivity of the thermocouple tube may cause the temperature differences over the length of the tube to decrease. This effect can be estimated from the heat balance for the amount of heat flowing to the tube from the gas and the amount escaping to colder places in the tube. It can be shown that At = t,., - ttubc = & = outer diameter of the tube 44 = inner diameter of the tube X = heat conductivity of the tube material k, = heat transfer coefficient to tube surface. When this formula is applied to the conditions present in the reactor under consideration, it is found that the maximum temperature difference amounts to O-6 “C, which may be neglected.

Thus it may be concluded that the resulting temperature differences are of no importance, and furthermore that external transport has no considerable influence on the reaction rate for temperatures below 450 “C. (c) The thermo-couple tube has a disturbing effect on the packing of the catalyst particles; i.e. there may be a larger gas velocity and consequently a lower conversion rate and a smaller temperature gradient around the tubes. This effect cannot be estimated accurately, but if it were serious it should be found that the temperature measured at the bottom of the bed did not agree with the temperature at the outlet of the reactor, as is actually the case. Moreover, the total temperature increase measured in the tube is in accordance with the temperature calculated on the basis of the total conversion in the gas, the heat of reaction and the specific heat of the gas-mixture. The temperature profiles given in Figs. 3 and 4 show that the different thermocouples do not indicate a homogeneous temperature in the catalyst layers. We are inclined to ascribe the observed differences to random inhomogeneities in catalytic activity or in the porosity of the bed.

379

P. MARS

11.3 Quantitative analysis of

the

temperature

From (4), (5) and (6) k can be calculated from the slope of the temperature profile and the reaction conditions. dT,ldx is determined from each temperature profile measured, and the mean value is used for the calculation. This calculation can be performed for every point of the profiles but the accuracy is only satisfactory in the bending point. The temperatures of the bending points Tb differ only to a small extent, and the catalytic activity calculated can be ascribed to the mean temperature. It is obvious that for the calculation of the first-order rate constant a correction must be introduced for internal diffusion. This correction is appreciable because of the large particle size. The effectivity factor was calculated as described in Part I. Secondly, allowance must be made for the fact that the sulphur content of the gas differs somewhat from the standard sulphur concentration used in laboratory experiments. With the help of the laboratory experiments represented in Fig. 2 the

projile in the first reactor

The rate constant k of the catalyst layer at a depth x in the bed can be calculated from the
+0x- Pco,)

&co = - --g

From the adiabatic character follows that in the stationary

of the reaction it state:

PCO,- PCO~_ T/ - T, PCO,-PCQ/

(5)

T/-T,

(assuming a constant specific heat of the gasmixture, independent of temperature and degree of conversion), so that:

dpmx - PCO,-PCOJ dx

T,-

Table 3.

T,

dTx * dx

(6)

Results of the calculation of the catalytic activitiesfrom data obtained about the reactors I and II in the factory

Reactor I

Effectiveness factor EtT

k

TB

(h-9

0

Residual activity due to sintering A

--

kcorrected (h-1) -

-

AI, at the beginning of a working period (Fig. 3)

401

4410

0.165

1.00

29000

AI, at the end of a working period (Fig. 4)

407

3750

0.250

0.48

33m

Reactor II

AIia BII AIIb >

at the beginning of a working period

Eflectiveness facror Eff

T.” Cc)

(h$

357 365 356

1440 1800 1750

380

0.355 0.290 0.320

keorrectea 01-l) 4300 6600 7410

Factors governing the behaviour of the adiabatic water-gas shift reactor

of k were corrected for deviations from the standard sulphur percentage. Finally there is the gradual sintering of the catalyst under reaction conditions. Of the two examples to be given below the first was obtained shortly after the reactor had been refilled. Here the catalyst must have approximately the original activity. In the second example the catalyst had to be replaced shortly after the measurements. In the laboratory it was found that samples drawn from the centre of the reactor had a residual activity of about 38 per cent under standard conditions. To be able to compare the results of all calculations on the original activity level, the calculated value of k was multiplied by l/0*38. Table 3 gives the results.

values

II.4

Calculations of k from

10s

the second reactor

As mentioned above, no vertical thermo-coupIe tubes are present in the second reactor. As the temperature gradient is only small, it was decided to calculate k directly from the decrease of the CO content in the whole reactor. The catalytic activity was calculated with the same formula used for the laboratory experiments: however, qeq was calculated at the end temperature Td, whereas the calculated value of k is ascribed to the mean temperature in the bed, Obviously, the above-described corrections for internal diffusion and variations in the sulphur content are also introduced here. As the temperatures are much lower than in the first reactor, the sintering effect was negligible. The results of these calculations are also given in Table 3.

II.5

(1) the intrinsic activity of the catalyst at various temperatures; this may decrease slowly with time, due to sintering; (2) the retarding effect of the diffusion of CO in the catalyst pores, which decreases the efficiency to values of 5CL5 per cent (at the low and at the high temperature side of the reactor respectively) ; (3) the reversible poisoning level, which depends on the actual sulphur content in the gas.

Discussion

In Fig. 5 the results of laboratory experiments and those of the calculations from technical data of the Grst and second reactor are pIotted together in one Arrhenius plot. It is clear that a good agreement exists. From this agreement we are inclined to conclude that both the laboratory experiments and the method of calculation trOm the technical data are substantially correct and that the main factors governing the activity of the catalyst in an adiabatic water-gas shift r#Etor asp 381

I.43

1.47 I-51

t55

I.59

Ial

l97

WI

Doe/T

Comparison of the catalytic activities. 5. 0 activity measured in the laboratory x activity calculated from data about technical reactors

FIG.

The results obtained so far have induced us to proceed also in the other direction, namely to try to calculate from laboratory data the temperature profiles (and consequently the hydrogen production in technical reactors working under different conditions). The results of these calculations will be discussed in the next section.

PART III.

CALCXJLATIONOF THE TEMPERATURE

PRomEs mu I~-IEZ HYDMXXN PRODUCTIONIN

m

REACTOR ON THEBASISOF LABORATORY DATA

Firstly, we will calculate the temperature profile in the first reactor, using laboratory data and the composition and temperature of the feed gas. These results will be compared with the actual measured profiles. Secondly, we will calculate the

P. MARS

influence of variation of a few factors on the performance of the reactor. III.1

Formulae used for the calculations

The rate of the chemical reaction at point x in the reactor is given by the following first-order expression : _

ydpco

= Eff A k_exp( - JWT,)(P~~~ - PCOJ (7)

1~

where V, = the linear velocity at S.T.P. (cm h-‘) PCO~ and PCO,, = the actual partial pressure and the partial pressure at equilibrium of CO at point x A =sintering factor, see below Defining Y by PCO, = pco, - y, the equilibrium condition reads [lo] :

~)(hOc- v) (PCOzr+ Y)(PHzr+ v)

Pco,,~PH~o,,=(PCOc-

Keg=

PC%, hea

(8)

(where i stands for the inlet composition), from which y and thus pco, can be calculated. The effectiveness Eff is given by the formulae: Eff =-$A-J-)(seeRef.

[4j)

In the.application of this formula, the temperature dependence of Q and of cp has been taken into account. The factor A, which accounts for the degree of sintering of the catalyst, is a function of temperature and age of the catalyst. As remarked earlier it appears from laboratory investigations that after a working period of many years the activity of the whole bed has decreased to a considerable extent. It was found that the accessible surface area decreases proportionally. For samples which had been used in the temperature range of 510-550 “C the loss of activity agrees well with the findings of H~OGSCHAGEN and ZWIETERING[I I] in laboratory experiments, but for samples which had been used at lower temperatures the actual loss of activity appeared to be larger than predicted by the formula given by these authors. Therefore, in estimating the residual activity at lower temperatures, after a working period of a few months, this activity was assumed to be lower than predicted by this formula. Fig. 6 shows the values of A found after several years from analysis of the discarded catalyst and the values of A after a few months as a function of temperature. 100

(9a) g

where

m = $A

k’/D,ff)‘iZ

0.60 0.40

(9b)

0.20

The effective diffusion constant Datc,02 was measured with oxygen at room temperature [4], and Dwr under reaction conditions was calculated from

Dcff

=

Dcff,oz J

$$$ co

assuming Knudsen diffusion. (Moz and MCO are the molecular weights of oxygen and carbon monoxide respectively.) With the aid of the heat balance, the temperature at point x can be calculated: Tx = T + e CP

. PC01

;

pcox

IW

360

380

400

420

440

TempMolum,

460

4~0

5~)

321~

540

*c

FIG.6. Decrease in activity of the catalyst by sintering. In formulae : a few months from the start of the working period:* A = 1 - (TX - 638)476 x 10” After

*In Part II we did not take into account the sinteringfor our calculation of k in tlw ‘YWW reactor at the bending point of the tempemture e From the small degree of sintering at this tanpMura (&NY’) it appears that the conclusions drawn them M not afWted.

382

I I ; i

Factors governing the behaviour of the adiabatic water-gas shift reactor

At the end of a working period A = 1 - (7” - 349)1*55 x 10-’ III.2

Results

Using all available information, as expressed by the functions given above, the values of T, and pcox can be calculated as a function of the bed depth X. This has been done with the help of a computer. It appears that steps of 1 cm were small enough to obtain the accuracy required. In Figs. 3 and 4 the calculated profiles have been drawn as solid lines. It appears that the calculated curves lie well within the profiles measured in the reactors, and the hydrogen production calculated also agrees with that found in practice. If we should not take into account the effect of sintering in the case indicated in Fig. 3, the calculated curve rises to excessively high temperatures (dotted line). Although a rough agreement between calculated and experimental temperature profiles is found, closer inspection shows that there is a considerable discrepancy. The slopes of the two calculated curves decrease too little at the higher temperatures. This fact requires further investigation.

III.3

440 Y 6 420 2 O400 k g

380 360 340 320 300

Fro. 7.

Influence of inlet temperature on the temperature profile.

from the original calculations that the effectiveness under normal conditions amounts to 50 per cent at the lower temperatures and decreases to 5 per cent at the higher temperatures. When the particle diameter is decreased to 0.7, 0.6 and 0.5 times the original respective values, a much

Influence of dierent factors on the performance of the reactor

The agreement found so far between calculated and experimental hydrogen production of the first reactor induced us to calculate the influence of different factors on the behaviour of this reactor, use being made of the formulae given. Fig. 7 shows the infiuence of the inlet temperature TI. As is already known from practical experience, small changes in Tg are of great influence on the value of Tf and on the amount of CO converted. (In this figure the final CO content present in the outlet gas after condensation of the steam is also given.) Fig. 8 shows the effect of increasing activity of the catalyst because of a lower sulphur content. A better removal of sulphur compounds from the feed gas increases the performance of the reactor considerably. Finally, the influence of the particle size of the catalyst was calculated, since it appeared

FIG. 8.

383

Influence of sulphur concentration on the behaviour of the converter (concentration of HIS in arbitrary units).

due to Dr. R. Nottrot who carried out the programming for the computer (Bull type y -E.T.). The author is much indebted to Dr. C. van Heerden and Mr. P. Zwietering for valuable discussions. NOTATION A = sintering factor Cp = specific heat of gas-stream mixture cal mol-1 “C-1 Den = effective diffusion constant in the tablet cm* se& E = activation energy cal mol-1 Eff = effectiveness coefficient k = first-order rate constant h= first-order rate constant, calculated on true tablet

z 0; 0

XC-’

keu = experimental effective rate constant &

sf+&-1 cm-a;:l

iji

%

P5 P

= partial pressure of component x = total pressure

Q R T

= reaction heat = gas constant = Temperature

larger conversion is obtained. (See Fig. 9.) These examples demonstrate that it is possible with the information available from laboratory experiments to estimate the influence of various reaction conditions on the performance of the reactor.

fi ys T V ‘lea 3

cm h-1 = gas w@iW, &ulatd at N-T-P= space velocity,gas volumeN.T.P./h.be.dvolume h-1 = depth in the reactor bed cm = heat conductivity CalSC- cm-’ “C-1 = degreeof conversion = degreeof conversionat equilibrium = particle diameter cm

Acknowlerlgment-The experiments in the laboratory have been carried out by Mr. J. G. H. Maessen. Thanks are

- - - i refers to the inlet of the reactor - - -f refers to the outlet of the reactor

Boilc4n

FIG. 9. Influence of particle size on the behaviour of the convertor ((Ppin arbitrary units).

giJtgz!E:::

SEX+

1

atm atm cal mol-1 I.986 cal mol-1 “C-1

REFERENCES

111 HOIJGEN0. A. and WATSONK. M. Chemical Process Principles Part III. John Wiley, New York 1949. PI WICKEE. Chem. Engw. Sci. 1958 8 61. 0. Acta Chem. Scad. 1958 12 1095. [31 BELLING 0. and VONKRUSENSTIERNA J. Industr. Engng. Chem., 1955 47 906. [41 HOOOSCHAGEN J. J. Chem. Phys. 1953 21 159. 151 BOKHOVEN C. and H~~QSCHAGEN Fl VAN HEERDBN C. and NOBELP. To be published. P. in The Structure and Properties of Porous Materials (Edited by EVERETT D. H. and STONEF. S.) p. 287. r71 ZWIETERINO Butte~orths, London 1958. VI ATWOODK., ARNOLDM. R. and APPELE. G. Zndustr. Engng. Chem. 1930 22 1091. H. and &EN&!KER G. J. Prakt. Chem. 1958 6 315. L91 HPXNZB rto1 GEYERE. W. and BRUGBE. A. Tables of Properties of Gases. Longmans, Green, London 1948. J. and ZWIETERINO P. J. Chem. Phys. 1953 21 2224. U11 HOOGSCHAGEN WI B~R~LINI P. Chem. Engng. Sci. 1958 9 135.

384