Fluid catalytic cracking: modelling of an industrial riser

~

AP PA LE IY D C AT L SS I A: GENERAL

ELSEVIER

Applied Catalysis A: General 138 (1996) 381-405

Fluid catalytic cracking: modelling of an industrial riser F. Van Landeghem b, D. Nevicato a I. Pitault a M. Forissier P. Turlier a C. Derouin a J.R. Bernard a 9

b,*

9

a G~nie Catalytique des R~acteurs de Raffinnage, ELF-CNRS, Centre de Recherches EIfSolaize, B.P. 22, F69360 Solaize, France b Laboratoire de G~nie des Proc~d~s Catalytique, CNRS, ESCPE-Lyon, B.P. 2077, 69616, Villeurbanne, France

Received 31 July 1995; accepted 10 November 1995

Abstract

The functioning of the riser of catalytic cracking units was simulated using a new kinetic model of the reactions and an accurate description of gas and catalyst hydrodynamics in the riser section. The kinetic model is based on a molecular description of cracking and hydrogen transfer reactions. It may be seen as an intermediate model between simple lumping of cuts and the ' single events' method. Chemical lumps, such as alkanes, alkenes, cycloalkanes, alkenic cycloalkanes and aromatics, are defined in different cuts. Kinetic constants are determined from experimental results obtained in a laboratory scale tubular reactor (M.A.T). The data on gas and solid flow in the riser are assessed by various experimental techniques: isokinetic sampling, Pitot probe, 3,-ray tomography and tracing in a cold set-up or in industrial risers. The model results were compared with industrial yields. Keywords: Cracking; FCC; Kinetic modelling; Hydrodynamics; Fluidized beds (circulating)

I. Introduction

A classical fluid catalytic cracking plant is shown in Fig. 1. The feedstock is injected at the bottom of the riser as a spray of liquid at about 250°C. At this point, the catalyst flows from the regenerator at about 750°C. The mixing of catalyst particles and feedstock droplets induces feedstock vaporization and * Corresponding author. 0926-860X/96/$15.00 © 1996 Elsevier Science B.V. All rights reserved SSDI 0 9 2 6 - 8 6 0 X ( 9 5 ) 0 0 3 0 9 - 6

382

F. Van Landeghera et al. / Applied Catalysis A: General 138 (1996) 381-405 ~f

zS,--7

-

_~cracking

effluents

Cyclones\~

Stripper.

• , ~ ~ 7

i!

_~ (

~'~<-;.

flue gas out

~ " ~ - Regenerator

Riser i

• ~::y~

x,

-

c o m b u s t i o n air

5:

~r ",--

Feedstock

Fig. 1. Scheme of the industrial FCC ( U O P process),

reaction start up. This provokes a high gas velocity (5-15 m / s ) which entrains catalyst particles upwards. The gases flow through the riser for about 2 s, the solid for about 4 s, during which time the reactions take place. At the riser exit, the catalyst contains about 1 wt.-% coke. It is separated from the gases in cyclones. The catalyst is then flushed with a counter current flow of steam in the stripper to minimize hydrocarbon entrainment to the regenerator. The coke on the catalyst is burned off by air in the regenerator. The mean catalyst recycle time is about l0 min. The coke combustion provides the thermal energy necessary to vaporize the feedstock and to compensate for the reaction endothermicity. This energy is transported by the hot catalyst. The outlet products are separated in a distillation unit. To give some orders of magnitude, about 45 wt.-% of the feedstock is converted to gasoline (distillation C5-C12), 20 wt.-% to liquefied petroleum gases (C3-C4), 5 wt.-% to fuel gases (H2, C1, C2), 15 wt.-% to light cycle oil (LCO, distillation 220-350°C) and 10 wt.-% to heavy cycle oil (HCO, distillation > 350°C). The remaining 5 wt.-% is converted to coke. The FCC unit plays a very important role in an oil refinery, because it converts heavy fractions (vacuum distillates or some vacuum resids) to gasoline, C 3 - C 4 cuts and petrochemicals. The units and their catalysts are in continuous evolution. They must adapt to market changes: gasoline yield a n d / o r quality maximization, petrochemicals production, conversion of residues, environmental requirements. The catalysts are extremely active so that contact times are short and catalysis is limited by heat and mass transfer. A rapid partial deactivation due to coking occurs; it must be accounted for. Hydrodynamics in circulating fluidized beds are not well understood and they modify the catalyst's perfor-

F. Van Landeghem et al./ Applied Catalysis A: General 138 (1996) 381-405

383

mance. The intention of the authors is to quantify these interferences occurring in this very complex reactor. They therefore used tools at three different scales: (i) the classical Micro Activity Test (1-5 g catalyst) for kinetic studies. It is done in a transient plug flow reactor which allows to set up a detailed kinetic model, including the quick deactivation by coke fouling. (ii) a cold set up (1500 kg catalyst) for hydrodynamic studies. Air flows at 4-7 m / s in the riser, entraining catalyst with a flux up to 300 kg m -2 s - i . It enables the flow regime, catalyst and gas flow properties, catalyst overall and local hold up to be determined. This latter factor is of prime importance in a catalytic reactor, although it is often ignored for this kind of bed. (iii) several industrial plants (catalyst inventories from 100 to 150 tons) where the possibility of introducing probes in working risers was managed. This enables to check bed hydrodynamic properties, and to measure the chemical lump yields along the bed. This paper gives a general overview of work on these fields and more details on the work of the authors.

2. Cracking kinetic model The cracking reactions are catalyzed by acid sites. The formation of carbonium and carbenium ions and their evolution controls the reaction course. The difficulty comes from the large number (several thousands) of different molecules. Moreover, each molecule can potentially take part in a large number of reactions. To complicate the situation further, deactivation occurs rapidly at the same time. 2.1. Choosing a method 2.1.1. Lumping in distillation cuts The first method to obtain a kinetic representation was to lump molecules in distillation cuts and to consider pseudo-chemical reactions between these lumps. Several authors studied how to lump various molecules [ 1-6]. The first lumping results were obtained by Weekman [7], but they were further developed by numerous authors [8-16]. The lumps are often the feedstock and the final products (gasoline, gases and coke). It is known that the kinetic order of cracking single molecules is 1. However it increases when cracking a mixture with a wide range of reactivities, like vacuum gas oils. This is because the most reactive molecules of the feedstock lump disappear f'trst and the remaining molecules have a lower kinetic constant when conversion increases. An order of 2 is therefore often used.

384

F. Van Landeghem et al./ Applied Catalysis A: General 138 (1996) 381-405

2.1.2. The 'single events' model The need of incorporating chemical information into cracking kinetic models appears very early: Weekman [17] distinguishes feedstock components (alkanes, alkyl chains, cycloalkanes and aromatic cycles). In this case the reactivity spectrum is narrow enough to come back to a cracking order of 1. John and Wojciechowski [18] took into account the complete gas analysis. More recently, kinetic models based upon reaction mechanism and elementary steps were developed. Liguras and Allen [ 19,20] simulate the feedstock with a great number of different molecules but this model does not account for reactions like condensation or hydrogen transfer. A model of chain reactions is proposed [21,22] for reactions of pure and light hydrocarbons. The most advanced method is the 'single-events' method, proposed by Froment and co-workers [23-26]. It permits a mechanistic description of catalytic cracking. It is based on the detailed knowledge of the mechanism of the various reactions involving carbenium ions. Several general chemical rules are used: the single event constants depend on the carbon type and not on the molar mass, the kinetic constant of the elementary step is proportional to the number of single events, thermodynamic constraints for isomerisation and protonation reactions, allow for the reduction of the number of unknown constants to remain at about 50. The method of determining them uses some key reactions of pure hydrocarbons. Nevertheless, the application of this 'single-events' method to catalytic cracking of industrial feedstocks (vacuum gas oil), is difficult because of analytical complexity and computational limitations. Moreover, hydrogen transfer, coke formation and coke fouling are not yet considered. Another method based upon molecular reactions, called structure-oriented lumping was recently proposed [27]. 2.1.3. An intermediate molecular model We propose a kinetic model giving chemical information, applicable to a riser reactor model, obtained with reasonable effort of analysis and computation. The important chemical reactions occurring during catalytic cracking were listed by Gates et al. [28]. Some of them are in Fig. 2. These reactions involve alkanes, alkenes, cycloalkanes or aromatics. These molecules are present in the various cuts (gases, gasoline, gas oil, and feedstock). Consequently it is suggested to define lumps based on the chemical functions in each traditional cut and to account for chemical reactions which have a significant importance when modelling results obtained in a micro reactor (see Section 2.3.2.) with different feedstocks and detailed analysis of gasoline and gases. The kinetic constants of chemical reactions between lumps are then no longer empirical but they represent a mean value of the kinetic constants of all possible reactions of lump components. As in Weekman's model [17], the kinetic orders are those of the pure molecules, i.e. 1 for cracking and 2 for hydrogen transfer reactions. This technique is less detailed than the 'single-events' method, but it is

F. Van Landeghera et al./ Applied Catalysis A: General 138 (1996) 381-405 1. alkanes cracking : CnH2n+ 2

>

CmH2m + CpH2p+ 2

)-

CmH2m + CpH2p

385

with n = m + p

2. alkenes cracking :

CnH2n

with n = m + p

3. !3-scission of aromatic alkyl chains :

ArCnH2n+l

>

Ar-CmH2m_ 1 + CpH2p+ 2

)'

CnH2n olefin

with n = m + p

4. cycloalkanes cracking:

CnH2n naphthene

5. hydrogen transfer -cycloalkane + alkenes - coke precursor + alkenes

6. isomerisation : - alkene

-~

> aromatic + alkanes > coke + alkanes

branched alkene

7. condensation reactions :

RICH ~

CHR 2

> ~

R

2 R1

Fig. 2. The main reactions in catalytic cracking [28].

efficient enough to model the cracking of various vacuum distillates and to obtain information on product quality. The kinetic model is simple enough to be included in a reactor model with mass transfer limitations.

2.1.4. The rate expressions and the catalyst deactivation The rate of formation, ri~ (mol m -3 s- 1), of the product i from reactant r, through the reaction j is:

ri,j=aj.kj.C r. Oj(1 - e) where aj is the stoichiometric coefficient, kj the kinetic constant, C r the reactant concentration in mol m -3, ~j the deactivation function. (1 - e) is the local volume fraction occupied by the catalyst in the reactor. The deactivation factor q~j is difficult to determine because every kinetic test provokes coke fouling. This is why it is often predefined as an exponential or a hyperbolic function of the time on stream, whose parameters are fitted to the experimental results, together with the kinetic constants. This approach may be misleading because time is not the physical cause of deactivation and it is then impossible to use the same function for reactors with different times on stream (i.e. microreactor and riser).

386

F. Van Landeghem et al./Applied Catalysis A: General 138 (1996) 381-405

The coke is the main cause of the deactivation. The catalyst coke content, therefore, is used as the variable for deactivation [29]. Guisnet and Magnoux [30] showed coke is produced in the zeolitic part of the catalyst. Under cracking conditions, it consists of polyaromatic molecules which are blocked in zeolite cages, occupying catalytic sites, and preventing diffusion. The authors tried to develop a method to determine experimentally a deactivation factor which can be correlated to the coke content of the catalyst [33].

2.2. Experimental 2.2.1. A laboratory scale reactor for measurement of kinetics Several laboratory devices were used for kinetic measurements in catalytic cracking ([17,13]). The tubular plug flow reactor of the microactivity test (MAT) is preferred for its simplicity. The MAT is a normalised reactor (ASTM D 3907-92, from the ASTM committee D32) used with a standard feedstock to compare cracking catalysts. The thermal transfer is satisfactory thanks to the annular shape of the catalyst bed. The feedstock is vaporised in a particular section of the reactor, the effluent recovery and analysis is well known. The mass balance is verified for each experiment. Of course, the pilot plant, which has been described in detail [16,31 ], is not used under the standard conditions. Catalyst mass, feedstock flow-rate and mass, total pressure, feedstock composition and gas dilution can be easily modified. The conversion field is between 40 to 99%. The contact between catalyst and reactants is not exactly the same in a plug flow reactor and in a circulating bed. In the first one, the products appearing at the beginning of the feedstock injection are in contact with the active catalyst without coke. They can initiate coke deposition. In the circulating bed, coke is built up in the riser bottom, where heavy product and feedstock concentrations are high. The contribution of light products to coke formation may be lower. This can induce differences in the rate of coke deposition, total coke yield, and the deactivation function. However, no fundamental difference was detected between the analysis of coke originating from the two types of reactor but overall coke content is higher in MAT experiments. 2.2.2. The FCC catalyst The FCC catalyst appears as a powder of nearly spherical particles with a mean diameter of 70 tzm. Each particle consists of an amorphous silica-alumina matrix and about 20 wt.-%, of Y zeolite crystallites measuring some/zm. When the catalyst is taken from the unit after regeneration, it contains several impurities (Ni, V from petroleum impurities) and its physical properties are different from those of the initial fresh catalyst. It is named 'an equilibrium catalyst'. Table 1 gives some data on used catalysts.

F. Van Landeghem et al./ Applied Catalysis A: General 138 (1996) 381-405

387

Table 1 Typical catalyst properties of an equilibrium catalyst [31]

Activity: Conversion (ASTM norm.) (%)

69

Physical properties: Specific area (m2/g) Pore volume (cm3/g) Average bulk density (kg/m 3)

98 0.33 870

Chemical analysis: AI203 (wt.-%) Na (wt.-%) Fe (wt.-%) Residual coke (wt.-%) V (ppm) Ni (ppm) R%O 3 (wt.-%)

42.80 0.28 0.61 0.05 1250 890 1.62

2.2.3. The feedstocks Usual feedstocks (vacuum distillate) consist of hydrocarbons containing 20 to 50 carbons. It is a mixture of alkanes, cycloalkanes, aromatics (with one or several cycles) and some molecules containing S, N, and metal atoms. The Table 2 gives data on the three feedstocks used in this study.

Table 2 Analysis of vacuum distillate feedstocks [31] Aramco

Montmirail

Nigeria

930.1 11,78 0.84 2.750 1090 250 0.605 20.8 494

902.8 12.1 0,45 0,282 800 312 0.570 13.15 432

940,9 11.59 0.6 0.320 1470 765 0.610 17.4 368

Chemical composition (mass spectrometry) Alkanes (wt.-%) 18.0 Cycloalkanes (wt.-%) 18.3 Aromatics (wt.-%) 51.7 Polar (wt.-%) 12.0

21.6 30.7 39.6 8.2

12.4 34.8 43.5 9.3

7~P simula~d distil~t~n (K) l0 w.% 50 w.% 90w.%

679 736 807

658 718 779

Density 288 K (kg/m ~) KUOP Conradson Carbon (wt.-%) Sulfur (wt.-%) Total nitrogen content (ppm) Basic nitrogen content (ppm) C/H NMR aromatic C number (%) Molar weight (g/tool)

671 735 805

F. Van Landeghera et al./Applied Catalysis A: General 138 (1996) 381-405

388

2.2.4. Chemical analysis of feedstock and effluents All kinetic studies require an accurate analysis of the feedstock and effluents. Direct measurements of the concentration of each lump by chemical analysis of the effluents is desirable to compare experimental and model results and, thus, to determine the kinetic constants. Analysis of reaction products was carried out in four steps [31 ]: 1. the gas phase was analyzed by gas chromatography (H.P. Plot column), 2. the liquid phase cuts were determined by chromatographic simulated distillation (non-polar SP 2100 H.P. column), 3. detailed gasoline and gasoil analysis was performed by a OV1 H.P. capillary column on a Hewlett Packard 5890 gas chromatograph, the identification of peaks must be done carefully, 4. the carbon content of the catalyst was measured after reaction with a Leco automatic device. It used infrared detection of CO 2 formed during coke combustion of the sample under oxygen. Analysis accuracy can be estimated from replicate experiments. The standard deviation is lower than 10% of the measured value.

2.3. Results 2.3.1. An experimental determination of the deactivation function As discussed above, a deactivation function depending on catalyst coke content is necessary. It is important to obtain direct measurement of this deactivation function [16,32-34] to avoid introduction of new unknown parameters. A f'trst step is to study whether the deactivation function may be the same for all reactions. The study of selectivity versus initial catalyst coke content [34] (Fig. 3) shows that only coke yield is largely modified. Thus, a particular deactivation function will be defined for the coking reactions. However, as coke yield is about 5 wt.-% in the industrial units, the coke formation may be neglected in a first step and all reactions are supposed to possess the same deactivation function versus catalyst coke content. This deactivation function can be applied to the overall reaction: Feedstock ~ m products

(1)

Practically, m varies with the conversion, but its value lies between 4 and 6 and is never very different from its mean value. The kinetic expression of the reaction (1) may be written: r = k[feedstock]" • q~(c) where c is the local coke content of the catalyst. By assuming the gas phase to

F. Van Landeghem et a l . / Applied Catalysis A: General 138 (1996) 381-405

389

Selectivity 09:

{I}

~~

{2}

0.07

•

o7

0.06 005 •

{3}

0 04~

003

0.02 0 01

J

0o0 O0

' 01~

" 012

01,

' 014

" 015

o16

" 017

01.

" Oig

"

,.0

Conversion Fig. 3. Selectivity for gasoline {-1}, gas {2}, coke {3} versus conversion, without initial coke content (variable catalyst mass) ( * ) and with initial coke content (0 to 1 wt.-% and constant catalyst mass) (Q), T = 480°C, feedstock flow = 0.05 g s - J and ASTM feedstock, [32].

be in plug flow, the following mass balance in an elementary volume, Sd z, of the reactor is obtained: dF~

Sdz where d F l is the elemental molar flow. The feedstock molar flow and feedstock concentration may be expressed versus the conversion X: F x = Flo(1

-- X)

and 1-X [feedstock]

=

(m-1)+l+d

P

RT

390

F. Van Landeghem et al. / Applied Catalysis A: General 138 (1996) 381-405

F10 is the known initial feedstock molar flow and d the diluent flow. Taking into account these expressions,

Flo( RT)" ( (m-1)X+ S --fi1-X

l+d)"

dX=kC19(c)dz

This last expression can be integrated along the reactor. The first term is a computable function f of the conversion X

f(X)=kf t" ~ ( c ) d z Z=0

When the catalyst is initially coked with a constant c o value along the reactor, the measured conversion is the kinetic constant will be multiplied by the deactivation function corresponding to the initial coke. The integral corresponds only to the coke build up during reaction c b.

Xc,

f(Xc)=k4)(Co)

(Cb)dZ =0

If the catalyst is used without initial coke, the conversion X o is measured and

f(Xo)=k

~(Cb)dZ =0

If it is supposed that the integral has the same value in both cases,

f( Xc) ~(Co)=f(Xo ) This ratio is also a computable function with the same parameters. From the measurements of conversion with and without various values of initial coke, it is possible to obtain an experimental estimation of the deactivation function versus initial catalyst coke content. This function may be used to represent the effect of local coke content on the deactivation function. The shape of the deactivation function versus catalyst coke content (Fig. 4) may be explained by two phenomena [34]: 1. a chemical effect of the coke which occupies sites in zeolite cages and which prevents them from reacting. 2. a diffusional effect which results from the fact that when a cage is occupied by coke the molecules must use other ways to go through the zeolite crystal. It is expressed by the increase of the tortuosity. The results suggest that the diffusion effect in the zeolite part of the catalyst is important when the catalyst is coked. The experimentally determined deactivation function takes into account an effectiveness factor due to coke content. A mathematical expression of • versus catalyst coke content, c, is useful for

F. Van Landeghem et al./Applied Catalysis A: General 138 (1996) 381-405

1

391

U -"o ...... A,

0,8

[3

~

06

"".0 0 A ~

"'...

"~

0.4

~',-.

"~-7 - . 0.2 o

o

I

I

b

I

0.5

1

1.5

2

25

% COKE

Fig. 4. Deactivation function for vacuum distillates, Aramco (. • • O . • -), Montmirail ( - - - i , - - - ) , ( - - [] - - ) , 530°C, l bar and feedstock flow = 0.05 g s - l, [34].

Nigeria

the simulation. The previous interpretation of the shape of the deactivation function suggests that d~

= Fq~ + E@(1 - q~) dc where F is related to the chemical effect (proportional to the remaining activity) and E is related to the diffusional effect (proportional to the lost activity). This hypothesis leads us to formulate the following expression: ~=

E+F

E + Fexp[(E +F)c] The values of the parameters E and F are obtained by adjustment with the experimental values of • versus c. Unexpectedly, the values of these parameters are dependent of the used feedstock (Table 3). The different values may be due to feedstock impurities (the highest value of F is observed with the feedstock containing the most nitrogen compounds) or to

Table 3 Parameters of the deactivation function for different vacuum distillates, (T = 530°C, P = 1 bar, feedstock flow = 0.02 g / s and NOVA D catalyst from Grace)

F E cm

Aramco

Nigeria

Montmirail

0.17+0.10 2.03 + 0.70 1.59%

0.43+0.10 1.01 + 0.30 1.39%

0.40+0.10 0.86 + 0.40 1.26%

392

F. Van Landeghem et al./ Applied Catalysis A: General 138 (1996) 381-405

the kinetics of coke formation (the highest value of E is observed with the most aromatic and the heaviest feedstock). Another unexpected result is the limiting value c m of catalyst coke content which is rapidly attained, and depends also on feedstock composition and reaction conditions. It suggests that coke distribution in the zeolite crystallite is probably not constant. A deactivation function versus (c m - c ) is chosen for coking reactions.

2.3.2. Selection of significant reactions and lumps Literature data are insufficient to develop a kinetic model under industrial conditions because the relative importance of reactions may change with reactant (pure or in a petroleum cut), catalysts (pure zeolites, fresh FCC catalyst, steamed or equilibrium catalyst) and observed conversion. The kinetics of coke formation may explain these discrepancies. Experimental results with industrial feedstocks and equilibrium catalysts are necessary. The experimental yields versus conversion obtained with various masses of catalyst are convenient to detect important reactions that may not be neglected for the kinetic model [31]. This is the case for the hydrogen transfer reaction which consumes alkenes and cycloalkanes to form alkanes and aromatics. The aim of the model is to obtain data on gasoline quality through its chemical composition (alkanes, alkenes, cycloalkanes and aromatics). These lumps must be defined in the gasoline cuts. As they result from the cracking of molecules of the same type in the heavier cuts, the same lumps must be defined in these cuts. The lumps taken into account are listed in Table 4. The reactions between lumps chosen for the model are listed in Table 5. 2.4. Kinetic constant determination For each reaction, a kinetic expression must be formulated [13]. Some supplementary parameters must be added, for instance to represent the probabilities of reaction between Og and Oe in the H transfer reactions. The kinetic order of condensation and condensation plus hydrogen transfer reactions is supposed to equal one in comparison with each reactant (global order is two). An important part of the work is the computation of the yields of the reactor from a set of kinetic constants [35]. The reactor is a non-stationary plug flow reactor due to catalyst deactivation. The last step is to find the best set of constants to minimise deviations between experimental and calculated values. It is a well known technique but it must be used with care to prevent local minimum and strongly coupled constants. A first constant determination was done. It shows that it is possible to model the catalytic cracking with the supposed reactions. Constant ranking follows from known observations concerning the cracking of alkanes and alkenes: k~ >

F. Van Landeghem et al. / Applied Catalysis A: General 138 (1996) 381-405

393

Table 4 List of lumps Cut

Hydrocarbon chemical nature

Symbol

Analysis

Feedstock, > 350°C, MW = 400 g / m o l

alkanes cycloalkanes with 1 to 3 cycles cycloalkanes with more than 3 cycles aromatics

Pc Ncl Nc2 Ac

a a a a

LCO, 215-350°C, MW = 200 g / m o l

alkanes alkenes alkenic cycloalkanes cycloalkanes aromatics

PI O1 NOI NI A1

b b b b b

Gasoline, C5-215°C, MW = 100 g/tool

alkanes alkenes alkenes cycloalkanes cycloalkanes aromatics

Pe Oe NOe Ne Ae

c d d c c

Gas, C3-C 4, MW = 50 g / m o l

alkanes alkenes

Pg Og

c c

Coke

c

Coke, MW = 400 g / m o l

a The feedstock lumps are quantified only in the fresh feedstock, only their sum (i e, the unconverted feedstock yield) is known in the reaction products. b Only the sum of the LCO lumps is measured. c These lumps are individually measured. d Only the sum of Oe + NOe is determined.

Yield 0.25~ I

l'aral]huc Gasol n

o 20~ ,

• .

o,ls~

o.o

~ ~

/

,

_ / ~

J J"

•

/ I

-

~ -

Arcunatic Uasoline [

,

/ I

//

Catalyst Mass (g) Fig. 5. Experimental (points) and calculated (lines) yields and conversion in the MAT reactor with one g of the Nigeria feedstock injected in 50 s at 530°C.

394

F. Van Landeghem et al./Applied Catalysis A: General 138 (1996) 381--405

Table 5 List of reactions between lumps fl-scission

Cyelisation H transfer

N1 Ne NOI NOe Ac

+ + + + +

Condensation and H transfer AI + 2Oe + A1 + 4Og + Ae + Oe + Ae + 2Og +

Pc PI Pe Ol Oe Ac AI Ac AI Ncl Ncl Nc2 NI N1 NOI

--} ---} ~ ~ --* ---> ~ ---} ---} ---} ---} ~ ---} ---} ~

PI Pe Pg 20e 20g A1 Ae AI Ae NOI NI NOI NOe Ne NOe

Ol

---}

NI

i (Oe or Og) i (Oe or Og) j (Oe or Og) j (Oe or Og) m (Oe or Og)

---} --* ---} --* ~

A1 Ae AI Ae Coke

+ + + + +

i (Pe or Pg) i (Pe or Pg) j (Pe or Pg) j (Pe or Pg) m (Pe or Pg)

(17) (18) (19) (20) (21)

k (Oe k (Oe k (Oe k (Oe

---} ---} --, ---}

Coke Coke AI AI

+ + + +

k (Pe k (Pe k (Pe k (Pe

(22) (23) (24) (25)

or or or or

Og) Og) Og) Og)

+ + +

Ol Oe Og

+ + + + + + + + + +

O1 Oe PI Pe PI O1 NI Pe Oe Oe

(1) (2) (3) (4) (5) (6) (7) (8) (9) (10) (1 I) (12) (13) (14) (15) (16)

or Pg) or Pg) or Pg) or Pg)

k 2 > k 3, k 4 ~> k 5, k 5 > k 3 and k 4 > k 2 as shown by the values of the ratios k l / k 3 -~ 100, k 2 / k 3 -~ 20, k a / k 3 -- 200. However the errors of the kinetic constants have not yet been determined and the physical meaning of the constants is not established. Nevertheless, the model may be considered as an empirical one, and used for reactor simulations. The comparison of experimental results from a laboratory reactor and those calculated from the model gives a good agreement as shown in Fig. 5, although some discrepancies appear on some lumps.

3. Riser reactor modelling Since the fluidized catalyst is entrained by cracked hydrocarbons, the hydrodynamics of both phages play a role in this type of reactor. Several factors must be experimentally measured to set a model in which kinetics can be inserted: - - the catalyst hold up must be known; its distribution within the bed may be

F. Van Landeghem et a l . / Applied Catalysis A: General 138 (1996) 381-405

395

Cs (kglm3) 200

1.50.

I00_~..., &

\

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10

Fig. 6. E x p e r i m e n t a l solid concenmation axial profile in the cold set-up, ug = 5.2 m s - l, 216 k g s - ~ m - ~ ( * ), 186 k g s - I m - 2 ( m ) , 156 k g s -1 m - 2 ( ~ ) a n d 124 kg s - I m - 2 ( A ) , [44].

uneven because the reactor is fluidized. The local hold up (1 - e) is related to the local catalyst flux F c and the local catalyst velocity Uc by the relation: apparent catalyst density = (1 - e ) Pc = - -

uc

where Pc is the catalyst particle density. Uc can be expressed by the product of the gas velocity by a catalyst/gas slip ratio, to be determined. the gas flow properties must be measured, finally the mixing properties of each phase must be also quantified to value deviations from the ideal plug flow reactor. These measurements are done in a cold hydrodynamic set up (scale about 1 / 3 ) and also in commercial plants when possible [36,37]. -

-

-

-

3.1. Catalyst concentration 3.1.1. Mean catalyst density along the riser The mean catalyst density in a riser cross section can be measured by the pressure loss in this section because acceleration and wall friction terms are negligible. This is checked by independent measurements done with a ),-ray tomograph (see Section 3.1.2.). A typical result is shown in Fig. 6 [38]. At every circulation rate, there is first a density decrease due to the progressive acceleration of the catalyst. However, after the first half of the riser, the density increases surprisingly. This is due to the riser ending with an elbow shaped as a

396

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// o~ ¸o e/' / / 100

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blind tee to protect the line against erosion by a permanent mattress of catalyst. This effect is not well understood, it depends on the elbow design.

3.1.2. Catalyst density versus radial position It is measured in the cold set up and in the plant by "y-ray tomography [39-43]. Fig. 7 gives an example of catalyst density versus radial position in the riser of the cold set-up. There is a concentrated catalyst annulus at the walls and a dilute core in the centre. This phenomenon always exists, but it seems to become more pronounced as the average density increases [45].

3.2. Catalyst flux and gas velocity in the riser Solid flux and gas velocity can be measured by using an isokinetic sampling probe [38,44]. In this device, the suction gas flow is tuned such as to cancel the pressure drop of the probe orifice to sample the phases without disturbing their flow. The measurements show that the radial profile of gas velocity is not flat as would be expected from the very large Reynolds number of the gas phase alone. The presence of the solid affects the gas turbulence. The profile looks like a laminar flow with a progressive decrease in the velocity from the centre of the riser to the walls (Fig. 8). When combining this information with that of Fig. 7, it appears that the riser can be considered as a plug flow reactor only in a first approximation. For a precise evaluation, one must consider that the local contact time is large close to the walls (high catalyst density, low gas velocity) so that

F. Van Landeghem et al./Applied Catalysis A: General 138 (1996) 381-405

397

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gas velocity

radial p r o f i l e in the c o l d s e t - u p , F s = 303 kg s - 1 m - 2, ug = 6.7 m s - ~ and

h=5m.

gasoline selectivity may drop because of high conversion; it is the inverse in the centre, with a low conversion. The radial profile of the solid flux has a parabolic shape. Combining flux and density allows to calculate the slip factor. The free fall velocity of an average catalyst particle is in the order of 0.2 m / s . This would lead to a slip factor very close to 1 since the gas velocity is in the range 5 - 2 0 m / s . It is not the case because particles interact. They may also form clusters with a larger fall velocity.

3.3. The radial dispersion coefficient of gases The phenomena described above would have no consequence if the gas mixed radially very quickly. It is thus necessary to estimate the rate of this mixing. A radial catalyst mixing may exist, but it can be neglected for kinetic reasons: we have seen that the maximum coke content on the catalyst is very rapidly reached. Thus the whole riser is working with about the same coked catalyst having the same kinetic properties, so that the mixing extent makes no difference. The radial mixing of the gas is measured in the set-up [38,44] by tracing gas with helium. It is isokinetically introduced in the riser centre at a certain elevation. Its concentration along the diameter is measured at two different downstream elevations. This allows us to determine a radial dispersion coefficient D r based on Fick's law. In the set-up, the obtained values are slightly higher (0.002 m 2 s -~) than those obtained with the gas alone. The radial

398

F. Van Landeghern et al./Applied Catalysis A: General 138 (1996) 381-405

"~

40

~

~

--

2

'/ 115

5

/

t/

0 0

5

/

l 10

f 15

20

time (s)

Fig. 9. Example of catalyst density results obtained by tomography in the industrial riser.

dispersion is a lumping coefficient of several phenomena (diffusion, turbulence, gas entrainment by particles). It is expected to be larger in a larger riser.

4. C o m m e r c i a l plant m e a s u r e m e n t s

4.1. Catalystflow properties The local concentration is measured on site by tomography. It exhibits the same profiles as those observed in the cold set up. This conf'trms that the core-annulus structure also exists in large beds (diameter ca. 1 m). The flow is also traced by injecting a pulse of a few grams of radioactive catalyst. It contains lanthanum which emits y-rays after irradiation by neutrons, so that its passage can be extemally detected (Fig. 9). The most relevant

F. Van Landeghem et al. / Applied Catalysis A: General 138 (1996) 381-405

399

information is the average velocity between detectors which allows to calculate the average density since the circulation rate is known. The results confirm the density profiles shown in Fig. 6.

4.2. Gas flow properties The gas is also traced with a chemically inert tracer (argon or krypton). This allows us to obtain the mean values of the velocity, density and molar weight between detectors. The catalyst/gas slip ratio averaged between the detectors is also determined. It is also possible to compute the dispersion in the axial direction. It is generally small so that no significant deviation from plug flow is observed. However, when carefully considering the output, it appears that an interpretation based on the radial velocity profile of Fig. 8 would be more realistic: a steep front due to the high gas velocity in the centre followed by a long tail representing the low gas velocity close to the wall, rather than the Gaussian shape of a weakly axially dispersed pulse.

5. Simulation of the riser

5.1. Principle and hypothesis The modelling strategy derived from results given in Section 3. and Section 4.: For the gas: axial dispersion or backmixing can be neglected, inlet radial velocity profile and radial dispersion coefficients must be introduced, the gas expansion due to cracking will be given by the rates of reactions. For the catalyst: axial dispersion can be ignored and radial dispersion is not sensitive, the core-annulus structure of the local density must be reproduced; this is done by using the flux profiles and the slip ratios observed experimentally and introduced through empirical correlations. This leads to a 2-dimensional axisymmetric riser model. Thermal effects are still neglected at this stage. When the material balance is written [46], the following differential equation is obtained for a gas lump i:

OVgCi D r a (rOCi) a----Z = -;- a--; j + R, Steam is used in the unit, mainly to improve the feed injection. This steam verifies the differential equation with a production term equal to zero. The inlet

400

F. Van Landeghem et al./ Applied Catalysis A: General 138 (1996) 381-405

feedstock concentration is assumed to be constant whatever the value of r. The used boundary conditions are at the riser wall:

(0C/)

7"F-F r=R = 0

and at the riser centre:

- ~ r r=0 = 0 The feedstock is supposed to vaporise instantaneously at the reactor entry. With the flux of the gaseous oil and its total concentration, C T, the initial velocity profile is determined, using the correlation developed in the cold model unit (Section 3.2.). Knowing the solid concentration, the reaction rates are deduced, applying the kinetic model. From the solid velocity, obtained from the gas velocity and by using the slip factor ratio (VJVg), the contact time and the new concentrations can be calculated in each elementary cell. At the outlet, conversion and yield are obtained.

5.2. Simulation results The simulation may be used to validate the kinetic and hydrodynamic models and also to predict the effects of reaction conditions. The main features of a gas-solid flow are well simulated in our model. These features for the solid

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Fig. 11. Computed axial profiles of yields and conversion. Comparison with industrial results. Conversion ( • ) , gasoline yield ([]), light cycle oil yield ( • ) , liquified petroleum gas yield (O), coke yield ( • ), fuel gas yield (zx). Same operating conditions as Fig. 10.

phase are an acceleration zone at the bottom of the riser, a core-annulus structure (a diluted core and a dense annulus) in the fully developed zone, and a reconcentration of the solid at the top of the riser due to the abrupt geometry of the exit. Fig. 10 shows the radial profiles of the solid concentration at 1, 5, 10, 20 and 30 m height. The decrease of the concentration at the exit of the riser near the wall is due to the intense cracking which occurs near the walls where the catalyst is very concentrated. This cracking provokes an increase in the gas velocity, which in turn accelerates the solid, whose concentration decreases, as the solid flux is considered to remain constant in our model. Axial profiles of conversion and yields underline the importance of the cracking reactions in the first meters of the riser (Fig. 11), where the catalyst is concentrated and not yet deactivated, and where the feedstock is highly reactive. The conversion used in Fig. 11 is the normalised one (sum of the gasoline, LPG, FG and coke yields). The effective conversion, where only the unconverted feedstock is considered, is about 80-85%. The LCO yield remains nearly constant after his initial sharp increase, while the gasoline yield increases monotonously between 10 and 25 m. At the top of the riser, the slope of the gasoline profile increases slightly, due to the increase of the catalyst concentration at the exit. The distribution of the yields for a paraffinic feedstock obtained with the model at the riser exit is the following: FG 3.2 wt.-%, LPG 14.0 wt.-%, gasoline 43.5 wt.-%, unconverted feedstock 16.5 wt.-%, coke 4.9 wt.-%. These results are in good agreement with the data obtained in various industrial plants (Fig. 11). The total coke content of the catalyst, which is near 1 wt.-%, is attained very rapidly. As our deactivation function only depends on this coke content, the activity of the catalyst is constant (nearly 80%) over a long length of the reactor (Fig. 12). Fig. 13 shows the effect of the riser height on the gasoline composition. It underlines the importance of the hydrogen transfer reactions which increase the

402

F. Van Landeghem et al./ Applied Catalysis A: General 138 (1996) 381--405

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10

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~

20

25

30

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alkane and aromatic yields to the detriment of alkene yields, particularly from 20 m upwards. A first result of simulation is to evaluate the effect of the core annulus flow in

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Fig. 13. Computed axial profiles of yields of the gasoline lumps. Alkanes ( I ) , aromatics (zx), alkenes ( • ) , cycloalkanes (O), alkenic cycloalkanes (0). Same operating conditions as Fig. 10.

F. Van Landeghem et al./ Applied Catalysis A: General 138 (1996) 381-405

403

Table 6 Comparison of yields and conversions obtained with different types of model and industrial results Plug flow model

Radial dispersion model Typical industrial results

Dr(m 2 s -~)

c~

10 - z

C/O

6.0

6.0

6.0

5.7

Conversion (wt.-%) LCO yield (wt.-%) Gazoline (wt.-%) LPG (wt.-%) FG (wt.-%) Coke (wt,-%)

66.7 18.5 44.7 13.9 3.2 4.8

63.9 17.5 42,1 13,7 3.2 4.8

60.2 17.6 39.7 12.6 3.2 4.8

66.12 20.2 43.33 13.88 4.09 4.99

10 _3

the riser on the conversion and yields. As the reactant mixing in the riser is not very significant, it may be considered as two reactors in parallel: one (the core) with a low catalyst density and a high gas velocity which has a rather small conversion. The other (the annulus), with high catalyst density and low gas velocity, is in an overcracking regime. The global gasoline selectivity is not optimal. The calculation [38] shows that it may be 2 - 5 wt.-% larger if the phases obey perfect plug flow (Table 6).

6. Concluding remarks This work is an example of satisfying simulation results obtained combining a kinetic model determined in the laboratory and hydrodynamic data measured in the unit and in a cold set-up. The order of magnitude of conversion and yield are obtained without the use of unmeasured adjustment coefficients. The model is also validated by various samples from an industrial riser which cannot be reported in this paper. The kinetic measurements obtained with various industrial feedstocks with operating conditions close to industrial ones (temperature, conversion) is a main reason for the good results. The necessary simplifications of the kinetic expressions do not introduce errors larger than the accuracy of possible measurements of effluent compositions. The hydrodynamic data used in the model are also sufficiently accurate but are very empirical. Significant research is being done around the world in this area. In particular, the feedstock injection zone remains difficult to study and to simulate. The model gives information on the consequences of riser hydrodynamics on yields and selectivities. It can also be applied to evaluate effects of reaction conditions and reactor dimensions, but more understanding is necessary (such as the temperature effect on reaction rates, or the effect of reactor dimensions on hydrodynamic data) and it is the subject of further research.

404

F. Van Landeghem et aL / A p p 6 e d Catalysis A: General 138 (1996) 381-405

Acknowledgements The authors acknowledge G. Wild and M. Guisnet for valuable advice and R. Barberet and D. Gadolet for experimental work.

References [1] [2] [3] [4] [5] [6] [7] [8] [9] [10] [11] [12] [13] [14] [15] [16] [17] [18] [19] [20] [21] [22] [23] [24] [25] [26] [27] [28] [29] [30] [31] [32] [33] [34] [35] [36] [37]

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[38] M.P. Martin, C. Derouin, P. Turlier, M. Forissier, G. Wild and J.R. Bernard, Chem. Eng. Sci., 47 (1992) 2319. [39] M. Azzi, P. Turlier, J.R. Bernard and L. Garnero, Powder Technol., 67 (1991) 27. [40] L. Desbat and P. Turlier, 13th IMACS World congress on computation and applied mathematics, 22-26 July 1991, Trinity College, Dublin, Ireland. [41] L. Desbat and P. Turlier, ECAPT 92, Manchester, 26-29 March 1992. [42] L. Desbat and P. Turlier, ECAPT 93, Karlsruhe, 25-27 March 1993. [43] P. Turlier, L. Desbat, P. and J.R. Bernard, Process Tomography 95, Implementation For Industrial Processes, 6-9 April 1995, Bergen (Norway). [44] M.P. Martin, P. Turlier, J.R. Bernard and G. Wild, Powder Technol., 70 (1992) 249. [45] W. Zhang, Y. Tung and F. Johnsson, Chem. Eng. Sc., 46 (1991) 3045. [46] J. Villermaux, G6nie de la r6action chimique, conception et fonctionnement des r6acteurs, Lavoisier, 1982, p. 63.