Catalyst dilution for improved performance of laboratory trickle-flow reactors

Chemrcal Engmeenng Saence Vol 35, pp 59-66 Pergamon Press Ltd , 1980, Pnnted III Great Brltaln

8

CATALYST DILUTTON OF LAE%ORATOT\Y

J

van

FOR IMPROVED TRICKLE-FLOW

Kllnken

and

R

H

PERFORMANCE REACTOF?

van

Dongen

Kon~nM1Jke/Shell-Laborator1um, Amsterdam (Shell Research B.V ) Amsterdam PO Box 3003, The Netherlands

ABSTRACT

The low linear llquld velocltles In laboratory bench-scale reactors can give rise to flow maldlst.rLbutlon and hence to lnefflclent use of the catalyst bed. For testing catalysts In their orlglnel size and shape at practical space velocltles under trlckleflow condltlons dllutlon of the catalyst bed with small rnert particles 1s advocated The great merit of this bed dllutlon technique 1s demonstrated from residence tune dlstrlbutlons determrned with the help of radIotracers under practical condltlons of deep denltrogenatlon of a vacuum dlstrllate feed As a result of the addltlon of small partzcles axial dlsperslon 1s reduced substantially, to the extent that 011 plug-flow can be assumed In virtually all practical circumstances The much higher llquld hold-up In the diluted beds leads to reproved wetting and utlllzatlon of the catalyst particles as was demonstrated for deep denltrogenatlon (-9 $ nitrogen removal) of a vacuum dlstlllate feedstock under practical condltlons

KEXWORDS Hydrocarbon residence

processing, trickle-flow tzme dlstrrbutlon

reactor,

catalyst

dllutlon,

axial

dlsperslon,

INTRODUCTION Trickle-flow reactors have found wldespread appllcatlon predominantly In the 011 Industry for processes such as hydrosulfurlzatlon and hydrocracklng of dlstlllate fractions In recent years very large trickle-flow reactor systems have been built for the removal of sulfur from residual 011 These commercial reactors are oneortwo orders of magnitude taller than practical laboratory reactors As a consequence, the superflclal velocltles of gas and llquld In the laboratory reactor are too low by a factor of lo-100 If one adheres to reallstlc space velocltles It 1s generally recognized (Satterfleld, 1975) that these lower velocltles may give rise to undesirable effects on the reactor performance The downward flow ofthellquld phase IS largely establlshed by gravity under normal trickle-flow condltlons, consequently the llquld tends to drain from the reactor and the amount of llwuld adhering to the catalyst particles In the lnterstltlal voids m the bed (dynamic hold-up) 1s relatively small at low feed rates WlJffels, Verloop and Zulderweg (1974) have observed that under these unfavourable condltlons the llquld runs preferentially In channels, leaving stagnant portlons of llquld In dead parts of the reactor These authors ratlonallze the resulting malcontact phenomena as a sltuat;tlon of mlnlIIRun total energy, taking into account the klnetlc energy of the flowing llqurd and the surface tensions In the three phase system Various studies have been performed In attempts to arrive at workable descrlptlons of the lmperfectlons of small laboratory trickle-flow reactors (Satterfleld, 1975) ConsIderable attention has been given (Van SwaalJ, 1967, Mears, 1971, Schwartz and Roberts, 1973) to the axial dlsperlon model extended to various degrees of complexity to account for exchange of reactants between dynazuc and static hold-up A different approach was chosen by Henry and GIlbert (1973), who proposed that catalyst utlllzatlon 1s related to the llquld hold-up, which 1s a function of llquld velocity Although this 59

60

Reactor

in prlnclple obJectlons Ity of the

A problem contacting of a size It should trickle-beds

Models

B-a

- Desrgn Studm

an accepted concept, Mears (1974) can be made against elaboration of correlations used by these authors

has polnted the Idea and

out that he also

a number questlons

of serious the valid-

1s that most of the correlations for backrmxlng or for hold-up and catalyst were obtalned on absorption towers packed with (nonporous) packing materials Apart from which and shape uncommon In 011 processing trickle-bed reactors be noted that, In general, llquld velocltles are considerably lower In than in absorption towers (Schwartz, Dudukovlc and Weger, 1976)

to correct for the nonldeal It thus appears to be very dlfflcult, of not lmposslble, behavlour of laboratory trickle-bed reactors. Against this background the present lnvestlgatlon was undertaken with the aim of ascertalnlng the merits of catalyst bed dllutlon as a means of remedying shortcomings of a small reactor rather than trying to Dllutlon of the catalyst bed with small inert particles 1s known to have quantify them a beneflclal effect on catalyst contacting efflclency and reduces backmlxlng, as has been mentloned by various >nvestlgators (Mears, 1971, Sedrlks and Kenney, 1973, De B?XrJn, 1976, Koros, 1976) It was consldered worthwhlle, therefore, to study the performance of a bench-scale trickle-bed reactor (91 cm bed height) under actual condotlons of vacuum dlstlllate hydroprocesslng Residence time dlstrlbutlons were measured for packed beds of catalyst beads and of nonporous glass beads both undiluted and These measurements provide lnformatlon on axial diluted with slllcon carblde powder dlsperslon and catalyst utlllzatlon which IS compared to the degree of denltrogenatlon achieved

RESIDENCE

TIME

DISTRIBUTION

MODEL

The residence time dlstrlbutlon In a "closed vessel" enters and leaves solely by plug-flow) 1s ldentlcal upstream pulse of some suItable tracer (Levensplel, response an E-curve, 1-t holds, by deflnltlon

.( The

mean

E(t)dt=l

residence

variance

time

t

z= The

u2

5

can

The total llquld follows llquld,

L

1s H

The

total

may

be

reactor

from

the

E-curve

according

to

defined

as

-1 Htot, mean

length

the volume fraction residence time since

and

v

the

superflclal

of

total

llquld

reactor

volume

velocity

Hence

occupied

VE

tot

by

(3)

tot=L

hold-up H

obtalned

(1)

hold-up from the

the

be

E(t)dt

a’=+&*E(t)dt

where

(defYned as a systemInwhlch fluid with the normalized response to an 1972) Calling such a normalized

can

be

consldered

to

be

composed

of

a dynamic

part

and

a

static

part

=H dyn +Hstat

An appropriate model for the residence time dlstrlbutlon In trickle-bed columns has He proposes axial dlsperslon In the mobile llquld been elaborated by Van SwaaJ (1967) In fact, the model 1s a phase with mass exchange between static and dynamic fluId comblnatlon of the crossflow and the dlsperslon models (Schwartz and Roberts, 1973) A well-deslgned laboratory trickle-bed reactor, being as long and as narrow as deemed but It ~111 feature enough staglng practicable, may not be a perfect plug-flow reactor, Ignoring all higher terms, we to allow a dramatic slmpllflcatlon of the mathematics arrive at the following expresslon for the variance

61

Catalystdllutlonforlmproved performance oflaboratory trickleflow reactors

E-4

The first The width of the E-curve thus appears to be the sum of two contrlbutlons Here Pe 1s the bed Peclet number defined term expresses the effect of ax~_al dlsperslon where u 1s the mean real llquld velocity, L the bed height and D the as Pe=uL/D, The parameter $J represents the fraction of the llquld coefflclent of axial dlsperslon. The second term In Equation (5) denotes the effect In the moblle phase, $=Hdyn/Htot The number of transfer stages 1s shown of mass transfer to and from the static llquld as Nt which 1s defined as ktAL/u where kt 1s the coefflclent of mass transfer and A the surface area between static and dynamic hold-up per unrt reactor volume

EXPERIMENTAL The experiments were carried out In a bench-scale reactor of 1 m length andemptycrosssectlon of about 2 5 cm* The reactor was fltted with a central thermowell The catalyst the remalmlng space at the top and bottom of the 91 cm in all cases, bed height was reactor being fllled with glass beads For bed dllutlon slllcon carblde powder with an average particle size of 0 2 nnn was applied In such a way that small portlons of catalyst and equal volume portions of dlluent were loaded one after another while IntermedIate vlbratlon allowed settlement of the bed The catalyst used was a commercial Nl-Mo/Al203 catalyst In the form of 1 5 mm spheres It had a compacted bulk density of 0 64 g/ml and a pore volume of 0 7 ml/g Before use the catalyst was sulflded In situ In a stream of H2/H2S The hydrocarbon feed In the experiments with this catalyst was a Kuwait vacuum dIstIllate (bolllng range with at 25 OC 0.9 kg/l, 630 ppmw nrtrogen) In the 5-95 % recovery, 305-510 OC, density reference experiments, this catalytically hydrotreated 011 was with 1 5 mm glass beads, used as feedstock so that as far as possible the llquld propertles were the same A survey of the reaction condltlons 1s given m Table 1 Note that all experiments were carried out at one llquld loading

TABLE

1

ReactIon

Condltlons

In

Llquld

compacted

volume

of

Gas

Once-through

Operation

MPa 12 5 1000 Nl/kg feed 0 125 kg m2s 1 9x lo- I m/s kg oll/m3 cat* 0 23 0 14 kg oll/m3 cat*

Hydrogen pressure Hydrogen rate Llquld load Superflclal llquld velocity Space velocity diluted beds Space velocity undiluted beds * basis

and

s s

catalyst

The residence time dlstrlbutlons were measured using a rsdlotr3cer, 1bC-labelled C32H66 This compound has an extrapolated normal bolllng point of 460 C, which falls in the tall of the bolllng range of the vacuum dIstIllate Immediately after lnJectlon of the tracer In the feed line, samples of the llquld effluent were collected over periods of one Mnute for one hour Values of the mean residence time and the variance were obtalned by the appropriate summations that may replace Equations (1) and (2)

AXIAL

DISPERSION

IN

THE

LIQUID

PHASE

Dllutlon of the 1 5 mm glass beads with an equal volume of 0 2 mm slllcon carbIde which 1s shlfted to longer particles results rn a residence time dlstrlbutlon (RTD) as can be seen In Fig 1 The shape of the E-curve does not seem to have changed times, the observed Increase In the mean residence time z 1s much According to Equation (31, the result of the Increased total llquld hold-up, going from 0 10 to 0 22 xnthe diluted bed (see Table 2) TABLE

2

Analysis

of

Residence

Time

Dlstrlbutlons

Undiluted Mean residence time, mln Variance, Total llquld hold-up,m3/m3 Peclet number, Bodensteln number,

79 0 18 0 10 11

0.019

Diluted 17 0 0 69 0

9 029 22 015

Reactor Models

62

NORMALIZED

- Destgn

Studies

B-8

CONCENTRATION 1-t

\ ‘i_ 0 Fig

1

RTD

for

glass

I

20 bead

packing

I

40

with

and

TIME,

without

60 mm

slllcon

carbIde

dlluent

It 1s not lmmedlately clear whether this Increased hold-up 1s due toan Increased static hold-up between the small dlluent particles or whether there 1s more 0x1 actually Independent measurement of static and dynamic hold-up underreactlon condltlons flowing was not possible In the equipment employed Nevertheless, dynamic and static hold-up values can be estimated If axlaldlspersI_on In the flowing 011 1s assumed to be negllgable, thus attrlbutlng the spread In the residence time solely to mass exchange In such a case between the flowing 011 and the static 011 adhered to the glass beads the moment of first breakthrough of the radlotracer In the effluent marks the uniform residence time In the flowing llquld The dynamic hold-up then can be calculated In the according to Equation (3) same way as the total hold-up, For the undiluted bed we thus find Hdyn'2' 0 04 and for the diluted bed Hdyn=O 17 Axla dlsperslon In the flowing llquld advances the lnltlal breakthrough of the tracer Particularly In the case of the undiluted glass beads unevenness of flow may be quite appreciable so that the estimated value of H d n may be much too low In conclusion, It would seem that by virtue of the small dlluen z particles the amount of oil. flowing In the reactor may be higher by a factor of four A markedly posltlve effect on catalyst wetting may therefore be expected From the total llquld hold-up (Table 2) and the lnltlal breakthrough of the tracer In case of the diluted bed of glass beads we calculate that the ratlo of dynamic llquld hold-up to total llquld hold-up $ 1s >O 77 The actual value of $ may be appreciably larger due to the occurrence of axial dlsperslon In the llquld phase, which results In to0 low VdUeS for Hdyn In view of the large value of $ In case of this diluted bed and the appreciable llquld hold-up In the pores of the catalyst (see below) it seems very reasonable to suppose In further calculations that Q= 1 In the case of glass beads, so that Equation (5) reduces to Pe= 2/a2 Table 4 shows that bed dllutlon leads to a sixfold Increase In the Peclet number This high value for Pe In the case of the diluted glass bed, equivalent to a cascade of 35 Ideal mixers, lndlcates that plug-flow of the llquld 1s approximated closely In these systems It was checked whether the use of the axial dlsperslon model 1s permlsslble In interpreting the data obtalned with the Good agreement was found between the experImenta residence tune dlstrlglass beads button curve and the one calculated from the model From the Peclet numbers the corresponding Bodensteln numbers can be calculated since diameter In a diluted bed the slllcon carblde Bo=udp/D where dp 1s the particle particles ~111 largely detenne the hydrodyn-cs so It seems reallstlc to put 1 5 mm and It 1s found to Boa 0 dp= 0 2 mm In that case With the undiluted bed dp= for both cases This figure 1s one order of magnitude smaller than the Bodensteln number corresponding to lamlnar flow In llquld-full systems It 1s also considerably smaller than the Bodensteln numbers reported for (absorber column) model systems (Sater and Levensplel, 1966, Hochman and Effron, 1969) Nevertheless, for want of a better descrlptron, we assume (Van Dongen and others, 1979) that Bo=O 02 for oil/ catalyst systems under trickle-flow hydrodenltrogenatlon condltlons, independent of particle size and llquld load

02

Catalyst dilution

B-8

CATALYST

for nnproved performanceoflaboratory

UTILIZATION

The presence of a porous catalyst has a large Influence on the residence the dlstrlbutxon peaks are broader and the mean residence times butlon In the corresponding cases with nonporous glass beads (Fxgs 2 and 3) NORMALIZED NORMALIZED

0

hold-up

CONCENTRATION

CATALYST

-i20

2

time dlstrlare hlnher than The static llquld

CONCENTRATION

‘,,,yRES

Fig

63

tnckle flow reactors

40

RTD for catalyst

In

the H

stat=

glass bead packings

catalyst

I 60 men

TIME,

and

pores

0

porous

follows

_c ,J’

Fig

CATALYST -. _S_PHERES . . . ._. --__ I 40

/’ 20

----__ TIME,

60 mm

RTD for glass bead and porous catalyst packxngs, both dxluted with slllcon carbxde powder

3

from

(H tot ) glass

(H totjcatalyst-

The static hold-up Itself has no absolute slgnxfxcance but should be compared to the total fraction of the reactor volume present In pores In the catalyst, F ore The ratlo Hstat/Fpore represents the fraction of the catalyst pores fllled with ox !? Table3 shows that by dllutlng the catalyst bed this pore volume utlllzatlon goes up by a factor of from 0.3 to 0 6 At first sight It would seem that an appreciable portlon of the two,

TABLE

3

Analysxs

of

Residence

Time

Dlstrxbutlons

Hydrocracklng

Undiluted Mean residence Variance,

tune,

mxn

Product nitrogen content, First order rate constant for nxtrogen removal,

number

of

mass

transfer

l-7 5 0 59

PFonW

20-250

kg/m3s

llquld hold-up, Iitot, total llquld hold-up In bed voids, llquxd hold-up In catalyst pores, catalyst pore volume, Fporea fraction of reactor volume, fractxon, Hstat/Fpore* p ore fllllng Nts

-

stages,

0.14-O

Diluted 31 7 0 091 14 50

I 40

m3/m3 m3/m3 m3/m3

0 22 0 10 0.12

0 0 0

40 22 18

m3/m3 m3/m3

0 0

45 27

o o

29 62

3

9

catalyst 1s still not yet fully utrllzed, particularly since pore fxlllng degrees as high as 0.9 have been reported In model studies (Schwartz, Dudukovlc and Weger, 1976) It should be realized, however, that the static hold-up Hstat was measured dynamlcally with the large C32 hydrocarbon molecules - to which the smallest pores may not be accessible The total pore volume Fpore, however, 1s measured by nitrogen capillary whxch fills all pores completely In addltlon, under the reactlon condocondensation, tlons, part of the pore volume may be occupied by gas.

64

Reactor

Mo&ds

-

Desrgn SW&es

B-S

More evidence of the Improved catalyst/o11 contact 1s provided by conslderzng the number of mass transfer stages, Nt Having establlshed appropriate values for Pe and by means of the reference experiments with glass beads, the *= (Htot) lass/(Htot)catr (5) value of Et can be calculated with the help of Equation As shown In Table 3, the number of transfer stages 1s trIpled by bed Since the diluted bed contains 35 % les? catalyst, the rate unit catalyst volume, kt_A, 1s even higher This Improvement must much larger wetted area A, assuming the mass transfer coefflclent determlned by pore dlffuslon and, therefore, not affected by the In the catalyst bed

dllutlng the catalyst of mass transfer per be attributed to a kt to be prlmarlly presence of a dlluent

With the undiluted bed the nitrogen contents of the reactor effluent appeared to fluctuate considerably (see Table F), which 1s not uncommon for this reaction system (Paraskos, Frayer and Shah, 1975) The corresponding first order rate constant for nitrogen removal varies between 0 14 and 0 50 kg/m3s (In our experience, concordant (1971), denltrogenatlon In a plug-flow reactor appears with the flndlngs or Mears first order In unconverted nitrogen) Dllutlon of the catalyst bed results In a llquld product with a very low nitrogen content which does not fluctuate The rate constant Increases to 1 40 kg/m3s which appears to be In line with the observed improvements In axial dlsperslon and catalyst wetting This good mutual agreement adds to our confldence In uslng the bed dllutlon technique to obviate the shortcomings of smaller scale reactors Having notlced the enormous contrlbutlon of bed dllutlon to rellablereactor performance the question arsses whether there 1s room for further improvement still As regards the trlckllng llquzd, zdeal plug-flow can be assumed to occur In diluted beds In virtually all cases Even In the present demanding case of 99 8 % nitrogen removal, Judging by the crlterlon formulated by Mears (1971), the observed Peclet number 1s so high that the hypothetIcally perfect plug-flow reactor needed to reach the same conversion level would be less than 10 % shorter than a real trickle-flow reactor

CONCLUSION Dllutlon of trickle-flow

the catalyst bed with reactor performance mlxlng

stages

1s

raised

to

particles

(1)

the number approximated,

(2)

the llquld hold-up 1s. Increased substantially, catalyst wetting The more effective wetting mass transfer stages

In practice wzll prove heavy feeds

of

small Inert In that a

level

1s

greatly

where which also

Improves

plug-flow results reflected

In

1s

the

laboratory

closely

a more effective In the number of

dllutlon of a bed of 1 5 mm particles with 0 2 mm slllcon carblde even for demanding cases such as deep hydrodenltrogenatlon adequate, In a 1 metre trickle-flow reactor

powder of

REFERENCES De

BTI-UIJ~,

Henry,

H

Hochman, Koros,

(1976)

A

C

, and

J.M R M

Levensplrl,

6th

J.B

, and

0

GIlbert

E,

(1976)

Intern.

Effron

4th

(1972)

(1973) (1969)

Intern

D

E

(1971).

Chem

Mears,

D

E

(1974)

Adv

Eng

Ind Ind

Symp

Chemical

Mears,

Congress

Chemistry

Eng Enp:

Reaction

Reaction Scl

Catalysis,

, 26, Ser

Chem Chem Enn

Englneerlng, 1361 , 133,

218

paper

B34,

, Process ~undls

Des , 8,

, preprxnts Wiley,

London Develop

72,

63_

1X-372, New

,

York,

Heidelberg Chap

9

328

Catalyst dhtlon

B-8

Frayer , J A 14, 315

Paraskos, J Develop

A

Sater,

and

V

E

,

Satterfleld.

C

Schwartz, Eng

W

Dongen, R SubmItted

Van

SwaalJ,

H

W

G

C

N

AIChe

W

Kenney

van In

F J 151

and 133,

4th

Eng

Eng

El;lk Eng

Chem

mndls

65

reactors

, 5,

, Process

Des

86

209

Ind

der Ind

tnckle-flow

Ena

Chem

(1976)

Elndhoven

Thesis,

Verloop Ser ,

21,

Chem

(1973)

Ind

Eng

(USA),

(1973)

of laboratory

(1975)

Shah

Ind

J

Roberts

(1967)

T

and E Weger Heidelberg

Bode,H , D for publlcatlon

J-B , J Chemistry

Y

performance

(1966)

Dudukovlc 1x-382,

and

and

Van

WlJffels, Adv

(1975)

N

J G

and

Levensplel

J G , M , preprints

Schwartz, 262 Sedrlks,

o

for Improved

Chem

Se1

, Process

28,

,

Symp

Chem

Des

ReactIon

Develop

,

12,

559

van Kllnken (1979) Develop , Process Des

and J Chem

Unlverslty

Zulderweg

Intern

of

Technology

(1974)

NOMENCLATURE A

statlc-dynarruc

Bo

udp/D,

D

coefflclent

Bodensteln

particle

d

InterfacIal

area

per

unit

reactor

volume,

m2/m

-

number,

of

axial

3

m2/s

dlsperslon,

m

size,

P normalized

E F F

pore void

H

tot

1 /s

concentration,

fraction

of

bed

volume

In

catalyst

fraction

of

bed

volume

In

lnterstltlal

total dynamic

llquld

hold-up,

llquld

fraction

hold-up,

m3 /m3

pores,

of

fraction

bed of

m3 /m3

voids,

m3 /m3

volume,

bed

m3 /m3

volume,

Hdyn H

stat

static

llquld

hold-up,

kt

coefflclent of statrc llquld,

L

catalyst

Nt

k+L/u,

Pe

uL/D,

t

time,

mass

bed

height,

number

of

Peclet

number,

fraction transfer

transfer

of between

stages,

bed

m3 /m3

volume,

dynemlc

and

Reactor Models - Design

66

mean

llquld

mean

real

residence llquld

superflclal H

dyr&ot

variance

’ of

dynsm~c response

B-8

S

time,

m/s

velocxty,

llquld

S&&es

velocxty, fraction E-curve,

m/s of

hqu1d,

m3/m3