Fluid Catalytic Cracking VIPreparation and Characterization of Catalysts, Proceedings of the 6th International Symposium in Fluid Cracking Catalysts (FCCs) [1 ed.] 9780444514738, 0444514732

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Fluid Catalytic Cracking VI Preparation and Characterization of Catalysts by Mario L. Occelli (Editor)

• ISBN: 0444514732 • Publisher: Elsevier Science & Technology Books • Pub. Date: July 2004

PREFACE

Since 1987, the Petroleum Division of the American Chemical Society (ACS) has sponsored at three year intervals an international symposium on fluid cracking catalysts (FCCs) technology. This volume collects the recent progress of this technology as reported in the 21 papers presented during the 226 th National Meeting of the ACS in New York, September 7-11,2003. Knowledge of the porous structure of FCCs is essential to the understanding of their catalytic properties. This volume contain chapters that describe in detail the use of density functional theory (DFT) methods to extract pore volume (PV) and surface area (SA) distribution data, from nitrogen sorption isotherms of commercially available FCCs. Results from the aforementioned DFT procedures indicate that traditional methods (such as the Langmuir and BET equation) currently in use can grossly underestimate SA and PV values in FCCs. In addition, the volume focuses on the use of modern spectroscopic techniques for the generation of detailed structural analysis required for the advancement of the science of FCC design. The use and importance of solid state nuclear magnetic resonance (NMR), microcalorimetry and atomic force microscopy (AFM) to the study of FCCs is described in several chapters. Scanning probe microscopies (such as AFM) can now almost routinely generate atomic-scale details of the surface topography of FCCs while providing information on surface configuration and porosity. Several chapters discuss strategies to control pollutant (sulphur or nitrogen oxides) emissions from a refinery FCCU and strategies to mitigate the deleterious effects of iron on FCC performance. Applied aspects of cracking operations are contained in a chapter describing on-line feedstock characterization and feedstock effects on FCC performance. Three of the chapters discuss strategies to improve olefin yields without compromising gasoline octanes. Advances in FCC preparation are contained in two chapters describing FCCs prepared by crystallizing the surface of an alumina matrix promoted with silica to generate mesoporous cavities coated with a layers of faujasite crystals and FCCs promoted by the gas phase deposition of alumina Finally I would like to thank colleagues everywhere for acting as technical referees. The views, comments and conclusions presented in this volume are those of the authors whom we sincerely thank for their time and effort in presenting their research at the symposium and for preparing the camera-ready manuscripts for this volume. The 7 th International Symposium on Advances in FCC Technology will be held during the 232 nd National Meeting of the ACS in San Francisco, CA., September 10-14, 2006. M. L. Occelli Atlanta, October 30, 2003

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 9 Elsevier B.V. All rights reserved.

An overview of physical adsorption methods for the characterization of finely divided and porous materials and their application to fluid cracking catalysts James P. Olivier Micromeritics Instrumem Corporation, Inc. Norcross GA 30093 USA Other than in degree of automation, basic gas adsorption measurement techniques have changed very little in the last half century, whereas our understanding of the adsorption process and the interpretation of adsorption data for the characterization of trmely divided or porous materials has advanced significantly in recent years. This contribution will attempt to review, somewhat critically, some of the older, more commonly used data reduction methods and present some of the newer insights gained from the application of simulation techniques such as density functional theory (DFT), Monte Carlo methods and Molecular Dynamics. 1. INTRODUCTION 1.1. The phenomenon The primary data of physical adsorption, expressed as the adsorption isotherm, is the measure of the equilibrium quantity of an adsorptive present at a gas/solid interface at a fixed temperature as a function of pressure. It follows that the magnitude of the quantity adsorbed at a particular pressure is a measure of the extent of that interface. If the data are normalized to a unit weight of sample as is customary, the isotherm comains desirable information about the specific surface area of the solid material. In addition to providing a measure of the extent of a surface, the adsorption isotherm conveys an abundance of information about the energetic heterogeneity and geometric topology or porosity of a solid substrate. As a source of information for the characterization of materials, the gas adsorption isotherm has much to recommend it: it is nondestructive, quite easily measured,

highly reproducible from laboratory to laboratory and provides a very sensitive indicator of surface properties. Using graphite at cryogenic temperature as an example, a 5% change in adsorptive potential would produce a twofold change in the pressure required to maintain the same amount of argon adsorbed [ 1]. The position of the adsorption equilibrium at a given temperature and pressure (a fixed chemical potemial) is totally governed by the strength of the adsorptive's interaction with the surface and any coexisting adsorbed film. This adsorptive potential is a measure of the net attractive energy between a solid surface and an adsorbed molecule. Adsorption is a spontaneous process; there is a decrease in the in internal energy (flee energy) of the system when a gas molecule becomes adsorbed. Thermodynamics dictates that this will result in the evolution of an exactly equivalent amount of heat, since no work need be done in the process.

1.2. The driving force For physical adsorption, the adsorption forces have theft origin chiefly in the London-type dispersion interactions (van der Waals forces) resulting from induced-dipole/induced-dipole and all higher multipolar attractions. These in turn depend on the size, polarizability and magnetic susceptibility of the interacting particles, which may be atoms or molecules. Additional significant contributions may come from Coulombic interactions that can arise from the direct interaction of polar adsorbates and adsorbents or be the result of induction forces brought about by the operation of a surface electric field or field gradient on induced or permanent dipoles or quadrupoles of resident molecules. The net potential due to these combined forces acting on an adsorbed molecule is generally short ranged, and is the summed effect from all nearby atoms. The contribution of the dispersion force to the total potential can be well approximated by the Lennard-Jones equation for pairwise interaction; ~b(s, e) = 4e[(~) 12 _ (~)6]

(1)

where f is the potential, e the depth of the potential well, s the molecular separation and ~ is the separation at which f = 0. The form of Eq. (1) for nitrogen is shown in Fig. 1. For an adsorptive molecule i at a surface, the total dispersion potential, U/LJ, is the sum of its pairwise interactions with each atom of the solid and all neighboring adsorbate atoms:

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Relative Separation (s/c) Fig. 1. Pairwise interaction potential of nitrogen according to Eq. (1).

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Relative Separation from Surface (s/c) Fig. 2. Molar adsorptive potential of nitrogen near a graphite surface.

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In Eq. (2), s~j is the distance from the/th adsorbed molecule to a neighboring atom j having an energy parameter 6 j. The case for a single molecule approaching a bare surface is illustrated in Fig. 2. The parameters chosen there are those represeming nitrogen and a graphite basal plane. The molar energy units give a better appreciation for the heat evolved. The dispersion forces account for virtually all the interaction in this example. The differences between Figs. 1 and 2 are subtle but significant. It is readily shown that over 90% of the value of U~ is provided by the surface atoms within 3 to 4 molecular diameters of the adsorbed molecule. It is clear that any irregularity in the local chemical composition, density or geometry of the surface will cause a significant variation in the adsorptive potential at that point. If the surface topography is locally reentrant, so as to constitute a pore of molecular dimensions, then the adsorptive potential can be much more than doubled. These effects are illustrated schematically in Fig. 3. Since the effects on the adsorption isotherm of surface compositional energetic heterogeneity and the pore topography are all manifested through this one adsorptive potemial, we must anticipate that there will be some difficulty in separating the two contributions when using adsorption data to characterize an adsorbem. 1.3. Some definitions It is conveniem when discussing porosity to have accepted classification terminology. The terms m icropore and macropore have been used with differem meanings across differem sciemific disciplines. To avoid confusion, the Imemational Union of Pure and Applied Chemistry (IUPAC) has advocated the foUowing def'mitions: pores of internal width less than 2 nm are designated micropores; pores with internal width between 2 nm and 50 nm are mesopores, those larger than 50 nm are called macropores. In general, for pores of non ideal shape, it is the smallest internal width that controls its adsorption behavior [10]. Note that these terms are not related to units of measurement but are based on the obvious Greek roots. The size ranges specified are not entirely arbitrary. The accepted definitions correspond quite well with certain natural process boundaries, i.e., micropores are in general created by either the inherent crystal structure of the material, as in zeolitic materials, or by a thermal decomposition process, as for carbonaceous chars and activated carbons. Generally, mesopores

Fig. 3. The Effect of Surface Topography on the Adsorptive Potential. A schematic representation of argon atoms (A- E) adsorbed on a rough surface. The largest shaded circles represent the region accounting for most of the Lennard-Jones 6-12 interaction energy. The small unfilled circles represent just the surface atoms of the adsorbent. It can be seen that even a slight ripple on the surface leads to a significant decrease (A) or increase (B) in the number of adsorbent atoms that can strongly interact with an adsorbate atom, compared to that on a flat area, (E). Adsorption at other locations will result in an energy change from about one-half (C) to more than double (D) that on a fiat surface such as at (E).

are the result of an agglomeration process, as exemplified by a silica or alumina xerogel or aerogel. As a result, the modes of the size distributions of the two types of pores are usually quite distinct, such that materials having both types of pores will show a bimodal distribution. Additionally, when studied by physical adsorption, micropores fill by a continuous volume filling process whereas mesopores fill by adsorption on the pore walls followed by adsorbate condensation. These differing processes call for different mathematical approaches in understanding and interpreting the adsorption data. Macropores much above 50 nm are simply more difficult technically to characterize by physical adsorption because the condensation pressure that measures their width approaches so closely to the vapor pressure of the bulk liquid adsorptive. The recent tendency to refer to all small pores as "nanopores" is really not helpful.

2. T H E O R E T I C A L TREATMENTS:

2.1 The classical methods The paradigm underlying the classical treatment of adsorption has its roots in the Gibbs' phase rule and bulk thermodynamics. Briefly, the picture is this: when a clean solid surface is introduced into a container of gas, a certain fraction of the gas molecules take up residence at the solid/gas interface. This spontaneous process is accompanied by the liberation of heat, a reduction in the gas pressure in the container and the appearance of a new phase at the surface. As in any classical system, the thermodynamic properties of this new system can be accounted for through the equations of state of the gas and the two-dimensional adsorbed phase. Both phases are considered homogeneous. Depending on conditions, the adsorbed phase can be treated as a two-dimensional lattice gas with or without adsorbate interaction (localized adsorption) or as a two-dimensional ideal or real gas (mobile adsorbed film). Once an equation of state has been chosen, an expression for the adsorption isotherm can be obtained through the Gibbs adsorption equation [1]. For example, a two-dimensional lattice gas without interaction yields the Langmuir isotherm. This paradigm works very well as long as molecular dimensions are negligible with respect to the system geometry, such that bulk thermodynamics can be applied. It is not applicable to adsorption in confined spaces of molecular dimensions, such as micropores. In the formative years of adsorption theory, it was customary to assume that a solid surface presented a set of equivalent sites for the attachment of an adsorbate molecule. Such was the idealized picture used by Langmuir [2] in developing his classic equation (Eq. 4). Even when the model was extended to include adsorbate- adsorbate interaction as was done by de Boer [3] in developing his two-dimensional real gas model for mobile adsorption, and by Fowler's statistical thermodynamic derivation [4] for the case of localized adsorption with interaction, the assumption of an energetically homogeneous, or homotattic [5] surface was retained, along with the assumption of monolayer adsorption. It remained for Brunauer, Emmett and Teller, [6] in 1938, who extended the Langrnuir concept to account for observed multilayer formation, to make the estimation of surface area from adsorption data a common technique. To aid in the discussions of data reduction techniques that follow, we will use the adsorption isotherm data obtained on a rather typical silica-alumina gel

500

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>>l, are the main source of the FCC microsphere surface porosity. Although the AFM is not an analytical tool, it can analyze the FCC surface for irregularities and defects that can be related to the manufacturing process used to prepare the catalyst. In fact, large surface discontinuities represented by micrometerscale trenches, valleys and craters appears whenever stacks of kaolin platelets are present. Moreover, the AFM can image surface debris on equilibrium FCC surfaces thus identifying the catalyst older fractions. It has long been recognized that carbon/conversion ratios measured during gas oil cracking depends on several parameters that include gas oil composition, reactor conditions and catalyst properties. The observation that at MAT conditions the catalyst of Figure 4D has a carbon/conversion ratio of~0.076 while the one in Figure 4A has a ratio of-~0.050 lend support to the possibility that micrometer-scale surface openings occlude and retain gas oil thus contributing to the catalyst coke production. AFM images recommend that carbon/ conversion ratios together with collision induced disintegration of FCC microsphere, fines generation, and greater catalyst make-up rates in a typical FCCU can be minimized by achieving a high degree of delamination of the clay during catalyst manufacturing so that the large trenches and crater-like surface openings shown in Figures 4-5 cannot form. The structure, composition and distribution of coke deposits in heterogeneous catalysts during hydrocarbon conversion reactions is and will probably remain a subject of controversy and debate. During gas oil conversion at MAT conditions with FCCs or PILCs, coke deposition results in the formation of surface fibrils containing chains of molecules shown in Figure 14 or in the formation of a layer geometrically similar to graphite shown in Figure 18.

102

Nanometer-scale contact mode AFM images of the (001) surface o f PILCs m Figures 16-17 clearly indicate that the steam instability of pillared montmorillonites can be safely attributed to the decomposition of the single silicate layer that separate the microporous structure while the greater structural stability of pillared rectorite can be attributed to the double silicate layers present. Thus in microporous pillared montmorillonites, steam collapse will occur irrespective of the hydrothermal stability of the pillars used.

Acknowledgement This work was supported in part by NATO collaborative grant CRG-971497 to MLO. Thanks are also due to B. Drake for valuable discussions and for providing AFM images during the initial part of this research.

REFERENCES [ 1] F. Mauge, J.C. Courcelle, Ph. Engelhard, P. Gallezot, and J. Grosmangin, New Developments in Zeolite Science and Technology, Y. Muricamiet al., Eds.; Elsevier, Amsterdam, 804 (1986) [2] P. Gallezot, B. Feron, M. Bourgogne, and Ph. Engelhard "Zeolites: Facts, Figures, Future", P.A. Jacob et al., Eds.; Elsevier, Amsterdam, 1281 (1989) [3] P. Gelin, and T. Des Courieres, Appl. Catal.72, 179 (1991) [4] R. A. Beyerlein, G.A. Tamborski, C.L. Marshall, B.L. Meyers, J.B. Hall, and B.J. Huggins Fluid Catalytic Cracking II: Concepts in Catalysts Design, M. L. Occelli and P. O'Connor Eds.; Am. Chem. Soc., Washington, DC., 109 (1991) [5] M. L. Occelli, D.C. Kowalczyk, and C.L. Kibby, Appl. Catal. 16,227(1985) [6] R. A. Beyerlein, C. Choi-Feng, J. B. Hall, B. J. Huggins, and G. J. Ray, "Fluid Catalytic Cracking III: Materials and Processes", M. L. Occelli and P.O'Connor Eds.; Am. Chem. Sot., Washington, DC, 81 (1994). [7] I. H. Musselman,P.E. Russel,R.T.Chang, M.G. Jamieson and L.C. Sawyer in" Proc. 12th Int. Congr. Electron Microscopy", Seattle,August 1990; W. Bailey Ed., San Francisco press, 866 (1990) [8] J. Binnig, H. Rohrer, C. Gerber, and E. Weibel Phys. Rev. Letters 50,120 (1983) [9] T. R. Albrecht, M.M. Dovek, C.A. Lang, P. Grutter, C.F. Quate, S.N.J ,Kuan, C. W. Frank, and R.F.W. Pease J.Appl. Phys. 64,1178 (1988). [ 10] S. Helveg, J.V. Lauritsen, E. Laegsgaard, I. Stensgaard, J.K. Norskov, B.S. Clausen, H. Topsoe and F. Besenbacher, Phys. Rev. Lett. 84, 5, 951-954 (2000) [ 11] M. L. Occelli, S. Gould and B. Drake in "Fluid Catalytic Cracking III: Materials and Processes," ACS Symp. Series, Vo.571, M.L. Occelli, P. O'Connor Eds.; 20 (1994) [ 12] M. L. OcceUi, S.A.C. Gould, Baldiraghi, S. Leoncini in "FluidCracking Catalysts" M. L.OcceUi and P. O'Connor Eds.; Marcel-Dekker, N.Y., N.Y.p.227 (1998).

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[13] H. Hartman, G. Sposito, A. Yang, S. Manne, S. Gould, and P. Hansma, P. Clays and Clay Miner. 38, 337-342 (1990). [14] F.J. Wicks, G.S. Henderson, and G.A. Vrdolijak, "Atomic and molecular scale imaging of layered and other mineral structures ; Scanning probe microscopy of Clay Minerals", CMS Workshop Lectures, Vol.7; Nagy, K.L., Blum, A.E., Eds.; 92-138 (1994). [15] M. L. Occelli, B. Drake, and S.A.C. Gould, J. Catal., 142, 337 (1993). [ 16] K. L. Nagy, A.E. Blum, "Scanning Probe Microscopy of Clay Minerals", CMS Workshop Lectures, Vol.7, The Clay Mineral Society, Boulder CO (1994). [ 17] G. A. Samorjai, M.X. Yang, J. Mol. Catal A Chem. 115, 389-403 (1997) [18] K. Fukui, H. Onishi, Y. Iwasawa, Phys. Rev Lett 79, 4202-4205 (1997) [ 19] A. Da Costa, Ch. Mathieu, Y. Barbaux, H. Poelman, G. Dalmai-Vennik, L. Fiermans, Surf. Sci 370, 339-344 (1997) [20] E.M Gaigneaux, Current Opinion in Solid State & Material Science, 3,343-353 (1998) [21 ] H. Lindgreen, J. Gamaes, P.L. Hansen, F. Besenbacher, E. Laegsgaard, P. Steinsgaard, S. Gould, and P. Hansma, Am. Miner., 76, 1218-1222 (1991). [22] F. Ohnesorge, G. Binnig, G., Science, 260, 1451-1456 (1993). [23] R. Giessell, R., Science, 267, 68 (1995) [24] M. L. Occelli, J. P Olivier, A. Petre and A. Auroux, J. Phys. Chem. B 107 (17): 4128-4136

(2003) [25] M. L. Occelli, H. Eckert, A. Wolker, M. Kalwei, A. Auroux, S.A.C. Gould; J. Catal. 196, 134-148 (2000). [26] M.L. Occelli in "Catalysts in Petroleum Refining and Petrochemical Industries" M. Absi-Halabi et al., Eds.; Studies in Surface Science and Catalysis, Elsevier, Amsterdam, 100, 27 (1995) [27] R.L. Smith, G.S. Rohrer, K.S. Lee, D-K. Seo, M-H. Whangbo, Surf. Sci., 367, 87-95 (1996) [28] M. L. Occelli, J. P. Olivier, and A. Auroux, J. Catal. 209, pp.385-393 (2002) [29] M.L.Occelli, Physicochemical Properties of Pillared Clay Catalysts. In "Keynotes in Energy Related Catalysis", S. Kaliaguin ed., Elsevier, Amsterdam, 101-137 (1989). [30] M.L. Occelli, A. Bertrand, S. Gould, J.M. Dominguez, Micro. and Macro. Mat 34, 2, 195-206 (2000). [31] R.E.Grim, "Clay Minerology", McGrow-Hill. N.Y., N.Y (1968) [32] S.Thomas, S., and M.L. Occelli, Clays and Clay Minerals, 48, 2 (2000). [33] R.J. Lussier, J.S. Magee,, and D.E.W.Vaughan, in " 7th Canadian Symposium on Catalysis," p. 112 (1980) [34] M.L. Occelli, I&EC Prod. Res. andDev. 3, 22, 4, pp.553-559 (1983). [35] J. Figueras, Cat. Rev., Sci. & Eng., 30(3),457-99 (1988) [36] J. Guan, E. Min, Z. Yu, H. Zheng and Y. Liang ; China-Japan-USA Symp. Hetero. Catal. Related to Energy Problems ; paper no.BO2C, p. 7-11 (1982). [37] J. Guan, E. Min and Z.Yu, US. Patent 4,757,040 (1987). [38] M.L. Occelli, Scientific Basis for the Preparation of Heterogeneous Catalysts ; Fifth Int. Symp. ; G. Poncelet, P.A. Jacobs, P.Grange and B. Delmond Eds. ; Elsevier, Amsterdam,287-299 (1991). [39] A. Wawkuschewski, H-J. Cantow, S.N. Magonov, Adv. Mater. 6, 6, 476 (1994). [40] J.E. Weisenhorn, J.E. MacDougal, S. Gould, D.S. Cox, W.S. Wise, P. Maivold, V.B. Elings, P.K. Hansma, and G.D. Stucky, Science, 247, 1330 (1990).

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[41] S. Thomas, J.A. Bertrand, M.L. Occelli, J.M. Stencel, S. Gould, Chem. Mat. Vol. 11, No. 4, pp. 1153-1164(1999) [42] M. L. Occelli, and S.A.C. Gould J. Catal., 198, 41-46 (2001). [43] M. L. Occelli, and J. Lester, Ind. Eng,.Chem. Prod. Res. Dev., 24, 27-32 (1985). [44] R. G. Haldemann and M.C. Botty, Jr. Phys. Chem, 63,489 (1959) [45] E. M. Gaigneauex, P. Ruiz, E.E. Wolf, B. Delmond, J. Catal. 172, 247-251 (1997) [46] S.A.C. Gould, K. Burke, P.K. Hansma, Physics Review B 40, 5363-5366 (1989).

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 9 Elsevier B.V. All rights reserved.

105

New Developments of NMR Spectroscopy Applied to Zeolite Catalysts Hubert Koller

Institut ~ r Physikalische Chemie, Schlossplatz 4/7, Westf~ilische Wilhelms-Universit~it MOnster, Germany

New and powerful solid-state NMR techniques have been developed in the last 10-15 years. This paper reports on some applications on the location and structure of acid sites in zeolites. The distribution and location of acid sites is one of the fundamental questions. By means of the 1 3 c - l i B and 13C-27A1 heteronuclear dipole interaction, an analysis of the spatial orientation of structure-directing agents towards the B and A1 atoms in the framework of a variety of zeolites (CIT-1, SSZ-24, Beta, SSZ-13, SSZ-16, ITQ-3, ITQ-4, ZSM12, Octadecasil) has been studied. This leads to interesting information about the mechanism of structure-direction and the location of heteroatoms implanted into the zeolite structure. Charge ordering between the negative framework charge (B or A1 sites) and quaternary ammonium cations was a general result observed by these dipolar NMR methods. ~B and 27A1 quadrupolar interactions in the activated acid forms of boron or aluminum zeolites are generally large due to a strong bond distortion around the heteroatom. These quadrupolar interactions are correlated with coordination symmetries and structural distortions. The dipolar interaction is also a useful probe to characterize the proton sites associated with distorted B or A1 framework atoms, and to investigate the nature of adsorbed species and their distances to the framework. A crucial advantage of the dipolar interaction compared to other NMR interactions is that it can be analyzed directly without the need of model assumptions.

106

1. INTRODUCTION Solid-State NMR has become an important method in zeolite research for the last twenty years. The state-of-the-art of the pioneering initial phase of highresolution NMR on zeolites in the 1980s, has been compiled in the book by G. Engelhardt and D. Michel. [1] Since publication of this monograph numerous new developments and applications of solid-state NMR techniques have emerged. A few of the more important ones are: multidimenional methods to correlate different NMR lines with each other can unravel connectivity patterns and solve ambiguities in assignments, [2] high-resolution techniques for quadrupolar nuclei [3-8] make a wealth of elements much more accessible, and powerful double-resonance methods enable the study of spatial neighborhoods and interatomic distances. [9-11] New in situ techniques allow the observation of a catalyst-reactant systems close to realistic working conditions. [12] Last but not least, the last 5-10 years have initiated the new age of multiple-quantum techniques which offer a wide range of applications. [13] All available NMR methods are aimed at the exploration of the important "NMR interactions" which contain the information wanted. Simply speaking, these interactions provide the framework of how the experimental and analytical strategies of NMR can be structured. These fundamental interactions are: a) The chemical shift interaction (chemical shift, 6~s) is still the most popular NMR parameter used. The isotropic chemical shift provides insight into the chemical surrounding of a nucleus, e.g. bonding partners, coordination number, and bond angles. If the chemical surrounding is anisotropic, then, in principle, the chemical shift interaction has an anisotropic component which can be measured and analyzed. b) The dipolar interaction is a magnetic interaction between two or more nuclei. This effect is transmitted through space which means that it is not restricted to chemical bond connectivities. The strength of the dipolar interaction is a function of the interatomic distance, and the dipole coupling constant, D, contains direct distance information. c) The quadrupolar interaction is an interaction between the quadrupole moment of nuclei with a spin quantum number of I > 1/2 local electric field gradient. It contains information about the local distribution. With the quadrupole coupling constant, Cq, information

nuclear and the chargecan be

107

obtained on the local symmetry: For ideal cubic, octahedral or tetrahedral coordination (high symmetry), Cq is zero. d) The indirect spin-spin coupling is a magnetic interaction between at least two NMR nuclei which is transmitted by bonding electrons. This interaction- sometimes also called J coupling- is only effective, when there is a chemical bond connectivity. This interaction is of lesser importance in so'lids, because the effect is usually too small to be resolved. This contribtution will concentrate on the magnetic dipole-dipole interaction and the electric quadrupole interaction which have contributed to a large extent to the NMR research of zeolitic acid sites in the last five years. This paper attempts to demonstrate the usefulness and versatility of the heteronuclear dipole interaction as a source of structural information on zeolites. The application of dipolar methods is ordered in three parts. The first part demonstrates, how dipolar methods can be used to obtain information on the location of heteroelements, B or A1, with respect to a guest molecule, in particular the structure directing agent. The results are important in light of the ultimate goal to implant active sites at specific positions in a zeolite. The second part addresses the local structure of the acid site in the activated zeolite. The difference between boron and aluminum in structural properties is outlined. The third part shows how the interaction between an acid site and a guest molecule is studied with dipolar methods. The example of methanol is used to demonstrate, how a protonated molecule can be distinguished from an unprotonated guest species by measuring interatomic distances.

2. EXPERIMENTAL The NMR experiments were carried out at a magnetic field of B0 = 11.7 T, where nothing else is indicated. Standard Bruker solid state NMR probeheads have been used. Sample rotation was achieved with evaporated N2 in order to prevent rehydration of the samples, where needed.

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3. LOCATION OF ACID SITES

Fig 1: Sketch of a structure-directing agent in a zeolite pore (example: carbon skeleton of methylsparteinium cation in SSZ24), and the translation of interatomic distance, r, into dipole coupling constant, D. ~to, y~, Ys, h, and rr are natural constants.

The location of active sites in zeolites is one of the challenges to an understanding of the catalytic properties. Due to the inherent long-range disorder of the distribution of these sites in most zeolites, it is difficult to locate them by diffraction methods. However, preferential siting may occur locally around the quaternary ammonium cations which fill the pores as structure directing agents (SDAs) in the synthesis of high-silica zeolites (Fig. 1). Such spatial information can be obtained via the heteronuclear dipolar interaction between 1H or 13C nuclei of the SDA and ~B or 27A1 nuclei in the zeolite framework. The dipole coupling constant of an isolated two-spin system, D, is easily translated into the interatomic distance, r (Fig. 1). In the past ten years, a set of sophisticated solid state NMR tools has been developed to probe local neighborhoods based on the heteronuclear dipole interaction in solids. This interaction is usually averaged out in conventional magic angle spinning (MAS) NMR experiments which is accompanied by the loss of the interesting information about the dipolar interaction. The basic idea of the rotational echo double resonance (REDOR) [9] or rotational echo adiabatic passage double resonance (REAPDOR) [ 14] methods is to reintroduce

109

this heteronuclear dipole interaction in the evolution period of a spin echo MAS NMR experiment, while retaining the high-resolution conditions by MAS during the data acquisition period. By this approach, it is possible to extract the dipolar information from the loss of echo intensity under high-resolution conditions. Here, the dipole interaction between ~H or ~3C nuclei of extraframework species (quaternary ammonium compounds as structure directing agents, methanol, and methoxy groups) and liB or 27A1 nuclei in the zeolite framework is measured. This approach allows one to obtain highly selective, and in some cases quantitative, information on internuclear distances and relative orientations between framework and extraframework species. [15]

Cll _~,/ C1 / N ~

C3 C4

C8//~C9 11~

~7/8~9 3 So

~__b)....~ c)

I

80

S - - So_S

~ |

60

I

I

40

20

5 / ppm

Fig. 2" 13C{1H}-llB CP REDOR experiment on zeolite CIT-1, Bo = 11.7 T

First, we demonstrate how these so-called double resonance methods can be utilized to measure the local neighborhood between I~B or 27A1 nuclei in the zeolite framework and the ~3C nuclei in the SDA. The results clearly show that the local charge distribution is not random. Rather, a preferred orientation of the charge center of the SDA towards the charge center in the zeolite framework (BO4/2 or A104/2 tetrahedra) exists. The first step in the REDOR experiment is to measure the line intensities (So) in a rotor synchronized 13C{IH} CPMAS spin echo NMR spectrum of the SDA in the zeolite (Fig. 2). Then, a second spectrum is measured, but now

110

additional dephasing pulses are applied (e.g. on the ~B resonance frequency) during the spin echo evolution time. The dephasing pulses cause the spin echo intensity to decrease (to the value S) for all those 13C nuclei which are influenced by a dipolar coupling to liB in the sample, i.e. which are in the neighborhood of boron in the framework. The larger the normalized difference intensity, (S0-S)/S0, for a given evolution time the shorter is the distance between 13C and ~B. Figure 2 shows the experimental results on zeolite CIT-1. [16] The REDOR effect, (S0-S)/S0 is most pronounced for the methyl groups of the quaternary ammonium group of the SDA (trimethylmyrtanylammonium cation). This observation proves that boron (negative charge center in the framework) is preferentially located near this charged group. Similar experiments have been carried out on other zeolites with boron in the framework. For aluminum zeolites, the REAPDOR technique has been applied. REDOR and REAPDOR data show that there exists an ordered orientation of the charged part of structure directing agent (SDA) towards the aluminium (or also boron) atoms in the zeolite. This result suggests that the choice of the SDA influences the distribution of acid sites in zeolites.

4. S T R U C T U R E OF THE ACID SITE 4.1 Boron

Zeolites containing boron instead of aluminum are known for their comparatively weak acidity. Tetrahedral, B [4], as well as trigonal boron, B TM, were found in boron zeolites by l~B MAS NMR spectroscopy, and a correlation with the nature of extraframework species has been suggested. [17,18] REDOR spectroscopy has been employed in order to show the selective association of B TMwith H + and of B [41 with Na + counterions. [1 9] The liB spin echo MAS NMR spectrum of dehydrated zeolite B-ZSM-5 shows two signals for B [3] and B [41. [20] The B [4] resonance is a sharp line at -3.7 ppm, while the B [3] resonance is characterized by a typical MAS powder pattern observed in the presence of second order quadrupolar interaction (Cq- 2.5 MHz, 1"1= 0.1, 8r - 9.7 ppm). The signal for B [41 exhibits a strong REDOR effect in the l lB{23Na} REDOR experiment, whereas B [31 shows only a small effect, indicating that only B t41 is associated with sodium cations. Similar observations have been made with other zeolite structures.

111

a ~ 2 9 3 ~ ~ . _

K

b)

623K

c) 293K 623K

~~_~82~K ~ ~ / ~ 2;

d

ppm

... .~

823K

' : d "i "2 '0 " IpPm"40;0

623K 823K

3s'oo

30bocm-'

Fig. 3: (a) lib MAS NMR spectra of H-B-Beta, (b) 1H MAS NMR spectra, (c) infrared spectra of zeolite H-B-Beta, dehydrated at the indicated temperatures; NMR spectra are measured at a magnetic field with B0 = 11.7 T.

Interestingly, in the hydrated state of zeolite Beta or ZSM-5, all boron ~ms in the zeolite framework show a tetrahedral coordination, and the trigonal won forms only upon dehydration. Different trigonal boron species have been postulated involving hydrolysis "a portion of the B-O-Si bridges. [21] Based on this work, we have embarked ~a systematic study of the incorporation of trigonal boron in the framework of olites. Here, we show the results of JH and ]~B solid state NMR spectroscopy conjunction with infrared spectroscopy on zeolite beta (Fig. 3). Different gonal boron species were found in the ~B MAS NMR spectra, and their ,undance changes with the dehydration temperature. A careful analysis of the ~B MAS NMR spectra shows that at least three fferent components with trigonal boron with different chemical shifts are esent (Fig. 4). The relative number densities of silanol groups, as found by ~H AS NMR and infrared spectroscopies, correlate with the number of boron ',nters which are partially hydrolyzed out of the framework, when different ',hydration temperatures are compared. The corresponding lines are assigned to partially hydrolyzed species, iOSi)~(OH)3_x (x-1,2,3). [22] The quadrupolar line with two maxima at ca. 8 Ld 1 ppm is assigned to B(OSi)3. The number of hydrolyzed bonds decreases ith increasing dehydration temperature which is illustrated by the decrease of e ~IB MAS NMR intensity between 10 and 20 ppm. The OH groups of the

112

hydrolyzed Si-O-B bridges are detected by ~H MAS NMR, and the condensation upon dehydration is confirmed by a decrease in the intensities in Fig. 3b. The infrared spectra (Figure 3c) show that hydrogen-bonded OH groups exist (3500 cm -~) at low dehydration temperatures which disappear at higher temperatures. This observation is consistent with the healing of the framework, because the presence of water molecules can be excluded based on the absence of bending vibrations for H20 in the samples studied here. The change in the ~H and liB MAS NMR as well as infrared spectra in Fig. 3 indicates a decrease in SiOH groups which means an increase in the number of S i-O-B bridges with increasing dehydration temperature. The three ~JB NMR lines, l-III, in Fig. 4 are assigned according to Table 1. Assuming equivalent chemical shift differences, the species B(OSi-)2OH would be expected around 13 ppm, but this line is obviously not resolved (or absent) in the liB MAN NMR spectra of zeolite H-B-Beta.

dehydration at 298 K

b)

e)

623 K

Z.

e) I

3; 2'0 1'0 ; 8/ppm

0

823 K

xl 3'0 2'0 1'0 ; 5/ppm

3; 2; 1}) ; 8/ppm

Fig. 4: Analysis of liB MAS NMR spectra of Fig. 3, B0 = 11.7 T.

113

Table 1" "B NMR line assignment line

III

assignment

6iso / ppm

B(OSi-)3

10.5 + 0.7

B(OSi-)OH2

15.4+ 0.5

B(OH)3

18.5 + 0.5

OH

OSi

I

I

HO/(. \OH SiO/ -

.-

20/

B(OSi) 4

OSi 10

d ; 3is o / ppm

B(OH)x(OSi)3_ x (x = 1,2) Fig. 5: liB chemical shift scale of boron in zeolites.

B(OH)3 shows an isotropic chemical shift of 19.9 ppm in solution. The chemical shift of component III (Table 1) is very close to this value. Therefore, it has been assigned to B(OH)3. Based on these results and on the work by others, [22,23] an ~IB chemical shift scale for trigonal boron in silicates is proposed (Fig. 5).

114 4.2 Aluminum

A Bronsted acid site is formed by a proton attached to a bridging oxygen atom between two tetrahedral atoms: silicon and aluminum (or boron etc.). This structure constitutes a formally three-coordinated oxygen atom. [24-26] The flexibility of boron coordination and incorporation in zeolite frameworks is intriguing, and one might conclude that aluminum should then have similar properties. However, the bonding properties can not be transferred from boron to aluminum. While quantum-chemical calculations had shown quite early [27] that trigonal boron exists in zeolites, trigonal aluminum is not expected. [8,28] Instead, aluminum atoms within acid sites have an extremely strained tetrahedral coordination with large 27A1 NMR quadrupole coupling constants of 11-18 MHz. [29] The Bronsted hydroxyl groups have been extensively studied by infrared and 1H NMR spectroscopies. [30-41] It often occurs that multiple ~H MAS NMR lines or infrared bands are observed which are difficult to assign. Some scientists, who contributed to the early developments of high-resolution solidstate NMR of zeolites, were proposing a correlation between ~H chemical shifts and acidities. It is now accepted that such a correlation is difficult to verify. Rather, hydrogen bonding is the dominating mechanism for the variation of proton chemical shifts. Hydrogen-bonded hydroxyl groups are observed in a broad range of 1H chemical shifts as well as IR frequencies. An important contribution has been made by Brunner et al., [42] who established a correlation between NMR and IR spectroscopy for such hydrogen-bonded hydroxyl groups: 81H / ppm = 57.1 - 0.0147(UoH / cm -1)

(1)

This empirical correlation between ~H NMR chemical shifts and IR frequencies performs well for a broad range of solids, and it allows one to assign IR frequencies to ~H NMR chemical shifts. This correlation was employed to analyze the various hydroxyl groups in a zeolite with IFR topology (Fig. 6). [43] Fig. 6 shows the 1H MAS NMR spectrum of an IFR zeolite with 1.64 A1/u.c., and the IR spectra of the same structure with various amounts of A1 per unit cell. The proton chemical shifts and IR frequencies are correlated with each other based on eqn. 1. It is not immediately clear which of the species corresponds to an acid site, and this question shall be tackled in the following part.

115

4.3 Merging Information from Dipolar and Quadrupolar Interaction The correlation in eqn. 1 and its application in Fig. 6 has an important consequence. In ~H NMR studies of zeolitic acid sites, it has been established that chemical shifts around 5-6 ppm can occur, if the bridging OH groups are hydrogen-bonded. However, a corresponding band around 3500 cm -~ in infrared spectra had not been interpreted as being due to acid sites, although correlations between IR frequencies and hydrogen bond strengths have been known for a long time. [44] The assignment of the band at 3488 cm -~ to an acid site is unusual and was also confirmed by pyridine adsorption experiments. Nevertheless, it seems desirable to confirm that the assignment of the corresponding ~H NMR chemical shifts at 6.3 and 5.2 ppm belong to acid sites. This can be done by merging the information from ~H and 27A1 NMR, because it is known that only acid sites show an extremely large 27A1 quadrupole coupling constant of larger than 11 MHz for A1 in a tetrahedral environment. [45] A favorable approach is a combination of dipolar and quadrupolar information which can be extracted from a single experiment. The 1H{27A1} dipolar interaction conveys the spatial information about the proximities, and the quadrupolar interaction of 27A1 provides insight into the structural distortion of aluminum. I IH MAS NMR ] 8~H/ ppm 6.3 2.8 5.2 ~ 3.9 2, ~/1.9 3.9 5.2 2.8 2.2 1.9 ~ 6.

IR Spectroscopy ]

UOH/ cml 3488 3629 3694

[

Ab15t

~

9

I

10

i

J

|

r

!

I

|

5

|

,

I

0

|

~

'

i

3488

I

r-~

A1/u.c. 1.64 0.92 0.68 0.44 0

3742

38OO i

3694

3700

3aoo

OOH/

35oo

340o

~300

c m -~

8 / ppm Fig. 6" Correlation between ~H NMR chemical shift and IR frequency of a H-zeolite with IFR topology, using eqn. 1.

116

The experimental solution of this strategy is realized by the 1H{27A1} REAPDOR technique. [11] It provides proton-aluminum distance information for each proton site, if 1H is the observed frequency. The 27A1 quadrupole interaction is probed by alternating the 27A| irradiation frequency systematically in various experiments, while observing the ~H frequency, i.e. the 27A1 frequency is mapped out by a set of REAPDOR experiments. By this method, the 27A1 NMR linewidth can be indirectly determined, giving the information about the structural distortion of the A1 environment. This approach has been applied earlier with another dipolar method, TRAPDOR, [10] but the REAPDOR technique has some advantages which are discussed in ref. 11. The result of such a 27A1 offset-dependent REAPDOR experiment on zeolite IFR is shown in Fig. 7. It is clear that the lines at 1.9 and 2.2 ppm do not show a REAPDOR effect. This is expected, because they are assigned to defect silanol groups. The frequency distribution of the 1H{Z7A1} REAPDOR effect, AS/S0, is larger the broader the 27A| NMR resonance line (or the larger the quadrupole coupling constant) of the coupling aluminum site is. The data in Fig. 7 clearly show that the proton sites with chemical shifts of 2.8, 3.9, 5.2, and 6.3 ppm are coupled with a aluminum sites with similar quadrupole coupling constant of 15 MHz which means similar structural distortions exists. Therefore, all these protons are spatially associated with aluminum in acid sites. The ~H NMR line at 2.8 ppm shows a lower maximum for on-resonance irradiation (at ca. 130 MHz) which means that this proton has a larger distance to the aluminum. On the other hand, the other three proton lines (3.9, 5.2, 6.3 ppm) have a similar maximum, and, accordingly, these protons have a similar H-A1 distance. Additionally, the ~H27A1 dipolar interaction is similar to the one observed in H-ZSM-5 for acid sites. [11] In conclusion, the three lines at 3.9, 5.2, and 6.3 ppm can be assigned unambiguously to Bronsted protons, while the line at 2.8 ppm is tentatively assigned to OH species which are close to acid sites, but have a larger H-A1 distance than within a Bronsted acid site. The different chemical shifts for acid sites arise from different strengths of hydrogen-bonding. These dipolar NMR methods provide powerful tools for an improved analysis of zeolitic acid sites, and line assignments can be bolstered. In addition, if the aluminum concentration is sufficiently low, so that multiple spin interactions can be excluded, a numerical analysis of REAPDOR data provides quantitativ H-A1 distance information. [11]

117

5. ADSORPTION COMPLEXES The aforementioned dipolar NMR techniques can also be applied to adsorption complexes. One of the most famous examples of the last ten years is the question as to whether methanol is protonated in a zeolite for a defined loading. [46] The two cases, a neutral methanol adsorption complex and a protonated molecule yielding a methoxonium ion, can be clearly distinguished by dipolar NMR methods, because they have different H-A1 distances. This problem has also been addressed using the REAPDOR technique. [47] Methanol-d6 was adsorbed in H-ZSM-5 with a 1"1 loading. Since the loaded sample was equilibrated at room temperature, the zeolite and methanol hydroxyl groups exchange their protons and deuterons quickly, and both sites yield ~H NMR signals. At 140 K two different proton chemical shifts at 4 and 14 ppm are observed, as also previously reported. [48] 1H{ZYA1} REAPDOR experiments at low termperatures with various evolution times (Fig. 8) show that the two different protons (SH] = 14 ppm, ~H2 - - 4 ppm) involved in hydrogen bonds between methanol and the zeolite framework have significantly different distances to aluminum which is consistent with the neutral adsorption complex.

0,8 --0-- 1.9 ppm 2.2 ppm

0,6

~

,,~

2.8 ppm

~v//"k\

3.9 ppm

/ . / / / ~ \~lll

5.2 pp~ --I-

Q 03

/~ ~ \

.

0,4

03 0,2

0,0

-0,2 122

124 27

126

128

130

132

134

136

AI irradiation frequency / MHz

Fig. 7" ]H-27AlREAPDOR experiment with 27A1frequency mapping of the IFR zeolite (of. Fig. 6).

118

Subsequent annealing produced several additional species which were analysed in terms of their distance to the aluminum of the acid site. lH{27A1} and 13C{27A1} REAPDOR experiments of these samples confirm the existence of surface methoxy groups in the annealed methanol loaded H-ZSM-5, albeit with a low number density. Fig. 9 shows the experimental data on a H-ZSM-5 zeolite which was first loaded (1"1 coverage) with CD3OD, then annealed at 200~ for 20 minutes, and finally evacuated for 60 minutes. The ~3C MAS NMR spectra show that the two major species (ca. 50 and 60 ppm) are removed by evacuation, and only a small portion of various species remains. 13C{27A1} REAPDOR experiments were carried out on the evacuated sample in order to probe the C-A1 distances quantitatively. The fits in Fig. 9 are numerical simulations which yield the distances as indicated in the Figure for the three lines at 57, 60, and 62 ppm. Only the line at 57 ppm can be assigned to surface methoxy groups, because its C-A1 distance of 3.1 A matches very well with the expected value from quantum-chemical calculations.

1,0 2.59 h

0,8 o

H1

r

0,6 0,4

o

~

9 H2

0,2 0,0

~,

|

i

0,0

0,2

0,4

|

i

|

|

0,6 0,8 1,0 1,2 evolution time / ms

|

1,4

Fig. 8" 1H-27A1REAPDORexperiment on a 1"1 loading of methanol (CD3OD) on H-ZSM-5. Numbers on the numerically fitted solid lines are H-A1 distances. This last example shows, how useful such analyses of dipolar interactions can be for a safer interpretation of NMR spectra. Discussions which are only based on chemical shift arguments always depend on an underlying model that provides the interpretation of the chemical shifts. In constrast, dipolar interactions do not require a model; they directly provide information on internuclear distances.

119

15 min at 200~ 57 15 min at 200~ 60 min vacuum 910 80

710 60

50 40 8 / ppm

310 20

10

1,0

9 57 ppm o 60 ppm

0,8

H"iY ~i~~ ~...

o

96

0,6

2

p

p

m

x

~

jt

4.4A

a
15% A1203) below 1270 K. Thus, for alumina-rich areas of the ECAT surface, and for the levels of Fe, Na, and Ca oxides we can expect to find there, the initial melting temperatures of such alumina-rich phases are expected to be hundreds of degrees higher than similar silica-rich phases. Every FCC catalyst microsphere is constituted of smaller particles of zeolite, matrix, clay, and the binder that holds everything together. As a result, the exterior surface of the particle is not homogeneous. Rather it has areas of

156

Table 4 [nitial melting temperatures for phase systems containing SiO2, A1203, FeO/Fe203, Na20 and/or CaO at ranges of concentrations which include those that can be found on the surface of FCC ECATs. Data adapted from refs. 10-17. System

SiO 2

SiO 2

AI203

(0/0)

(%)

(%)

First Liquid Temperature

(%)

(K) 1986

100

2323

Balance

AI203-Na20

SiO2-CaO

(%)

CaO

100

AI203 SiO2-Na20

FeO/Fe203 Na20

Balance

0*-25

1053

0*-5.5

2273

5.5-9.2

1853

>9.2

1683

Balance

AI203-CaO

Balance

0*-36

1709

0*-8.4

2123

SiO2-FeO

Balance

0*-62

1453

SiO2-FeO- Fe203

Balance

0*-85

1728

0"-35

2023

AI203-FeO SiO2-FeO-Na20 Na20-AI203-SiO2Fe203

Balance

0*-57*

25-0",t

40%) are those using the CD-ALFA conditions with iron added.

171

t a b l e 4. Deactivated catalysts properties Cat. A CD SA MSA deac. type cycles [m2/g] [m2/g] .... Fresh catalyst 272 72 Steaming 795~ 5 hrs 177 44 Steaming 788~ 20 hrs 155 36 CD-ALFA (no metals)* 200 206 55 CD-ALFA (Fe, V, Na) 200 136 31 CD-ALFA (Fe, V) 200 144 34 C D - A L F A (Fe, V, Ni, Na) 200 179 42 CD-ALFA (Fe) 200 200 52 CD-ALFA (V only)* 200 179 42 Conv. CD (Fe, V) 90 108 22 Conv. CD (Fe, V, Na) 90 82 18 Conv. CD (V, Ni) 45 150 33

Cat. B deac. type Fresh catalyst Steaming 795~ 5 hrs Steaming 788~ 20 hrs CD-ALFA (Fe) CD-ALFA (Fe, V, Na) CD-ALFA (Fe, V) CD-ALFA (Fe, V, Ni, Na) CD-ALFA (Fe) CD-ALFA (V only)* Conv. CD (Fe, Ni) * CD-ALFA

MiPV [cc/g] 0.093 0.063 0.056 0.070 0.049 0.051 0.064 0.069 0.064 0.040 0.030 0.054

CD SA MSA MiPV cycles [m2/g] [m2/g] [cc/g] 200 200 200 200 200 200 90

243 160 147 163 127 136 149 163 152 119

129 94 93 97 79 84 90 97 94 81

0.053 0.030 0.025 0.030 0.022 0.024 0.027 0.030 0.027 0.017

AAI Added Na20 [wt.%] 2.9 . 5.1 . 6.0 . 4.7 . 0.0 0.73 0.2 0.9 0.33 0.3 4.6 2.4 2.7 0.38 4.0 -

Added Added Added Fe V Ni [wt.%] [wt.%] [wt.%] . . . . . . . . . . . . 0.97 7403 1.13 8654 0.69 4541 972 0.96 6117 0.99 7794 0.94 6881 4671 1000

AAI Added Na20 [wt.%] 16.1 . 20.4 . 23.1 . 15.8 5.5 0.73 9.3 8.4 0.41 15.8 12.8 12.9 -

Added Added Added Fe V Ni [wt.%] [wt.%] [wt.%] . . . . . . . . . 0.97 1.16 6604 1.04 6695 0.60 4543 933 0.97 6022 0.57 853

steaming conditions but with no iron added

T h e c o n v e r s i o n s f o r e a c h d e a c t i v a t e d c a t a l y s t c l o s e l y t r a c k t h e total s u r f a c e a r e a , i n d e p e n d e n t o f A A I ( F i g . 5.) H o w e v e r , the L C O / b o t t o m s r a t i o d e m o n s t r a t e s a c l e a r c o r r e l a t i o n w i t h A A I , a n d is c o n s i s t e n t w i t h t h e e - c a t b e h a v i o r o b s e r v e d in the past.

172

Figure 4. AAI retentions (AAIfresh/AAIdeac)for various deactivations of Catalysts A and B. The impact of AAI on FST performance testing is illustrated in Fig. 6. This data set includes e-cats taken from over 80 refineries and encompasses all types of FCC technologies currently on the market. The full range of typical e-cat surface areas and metals levels are represented. Given the diversity of these e-cats, the relationship between LCO/bottoms ratio and AAI is particularly noteworthy. In fact, statistical analysis for this set confirms that AAI is the most influential measurement for predicting LCO/bottoms ratio in the FST from factors that include meso- and microporous surface area, total pore volume, catalyst density, and individual and collective metal levels. Fig. 7 shows the relationship between the LCO/bottoms ratio and AAI for the two deactivated catalyst sets (in Fig. 4 and Table 4.) The LCO/bottoms ratio qualitatively responds to these lab-induced AAI decreases in the same manner as observed for equilibrium catalysts (this is discussed in more detail below.) This suggests that the AAI reductions achieved in the lab using the CD-ALFA method are indeed related to the changes that occur in the commercial FCCU.

173

Figure 5. Conversion yield vs. deactivated surface area for Catalysts A and B. It should be noted that the LCO/bottoms ratio for the deactivated catalyst resulting from Catalyst A treated by CD-ALFA with iron alone (no added vanadium) does not appear to conform to the rest of the data. While this sample exhibits a significantly attenuated AAI (0.3), the lowered accessibility does not seem to be reflected in the performance testing. Also noteworthy, however, is that although the accessibility has dropped dramatically for this sample, the total surface area falls well above the range for typical e-cats, and is certainly much higher than all the other deactivated samples in this set. This, along with the absence of V in this deactivation, suggests that there are additional effects of surface area and metals interactions that have yet to be explained. For this data set, only catalysts deactivated with both Fe and V demonstrate significant AAI attenuation and the expected LCO/bottoms effects consistent with the e-cat testing.

173

Figure 5. Conversion yield vs. deactivated surface area for Catalysts A and B. It should be noted that the LCO/bottoms ratio for the deactivated catalyst resulting from Catalyst A treated by CD-ALFA with iron alone (no added vanadium) does not appear to conform to the rest of the data. While this sample exhibits a significantly attenuated AAI (0.3), the lowered accessibility does not seem to be reflected in the performance testing. Also noteworthy, however, is that although the accessibility has dropped dramatically for this sample, the total surface area falls well above the range for typical e-cats, and is certainly much higher than all the other deactivated samples in this set. This, along with the absence of V in this deactivation, suggests that there are additional effects of surface area and metals interactions that have yet to be explained. For this data set, only catalysts deactivated with both Fe and V demonstrate significant AAI attenuation and the expected LCO/bottoms effects consistent with the e-cat testing.

175

Figure 7.

LCO/bottoms vs. A A I at 70% conv. for deactivated catalysts sets A and B.

The FST features long catalyst-feed contact times relative to those of typical FCCU operations and this tends to damp out accessibility related performance effects. Nevertheless, the impact of AAI on LC0/bottoms yields in the FST is clear, both for the e-cats and the lab-deactivated samples studied here. If these two deactivated catalysts could be compared in an actual FCCU, the shorter contact time would lead to more dramatic performance differences. In fact, these differences would be expected to extend to the overall conversion and absolute bottoms yields, as has consistently been observed for comparable AAI differences in commercial data in a variety of applications [ 1-8]. 4. CONCLUSIONS FCC's undergo specific types of chemical and physical changes as a result of iron contamination, leading to decreased accessibility and deterioration in catalyst performance, particularly with regard to bottoms cracking. This decrease in accessibility is not simulated by most routine lab-scale deactivation methods, including Mitchell impregnation (along with derivative methods, such as Cyclic Propylene Steaming) and conventional CD.

176

The CD-ALFA method, which employs steaming conditions more commensurate with actual FCCU regenerators, allows significant AAI decreases to be induced as a function of added iron and vanadium, simulating the behavior observed in commercial FCCU's. Furthermore, these accessibility drops achieved in the lab impact the catalyst performance testing results in a manner consistent with that observed for commercial e-cats. That the absolute deactivated AAI level has a significant impact on FCCU performance has been amply demonstrated. This is reflected in the lab-scale testing reported here by the dependence of the LCO/bottoms ratio on AAI at the lower end of the accessibility scale. Because high accessibility catalysts (those with flesh AAI > 5) can undergo significant accessibility losses as a result of iron contamination, and yet still retain relatively high AAI's, they are more likely to allow the FCCU to operate in the regime above the critical accessibility threshold and are thus eminently suitable for high-iron, resid applications. Furthermore, the catalysts with the very highest fresh AAI's (>12) have demonstrated the best sustained performance in the face of elevated iron levels. REFERENCES [ 1] M. C. J. Hodgson, C. K. Looi, S. J. Yanik, Proc. Akzo Nobel Catalyst Symposium, Noordwijk, The Netherlands, June 1998. [2] E. Rautiainen and P. van Krugten, Catalyst Courier 40 (2000). [3] P. Imhof, E. Rautiainen, and K.Y. Yung, Catalyst Courier, 48, (2002). [4] P. A. Lane and S. J. Yanik, Catalyst Courier 41 (2000). [5] S. Foskett and E. Rautiainen, Hydrocarbon Processing (2001) 71. [6] A. K. Hakuli, P. Imhof, and C. W. Kuehler, Proc. Akzo Nobel Catalyst Symposium, Noordwijk, The Netherlands, June 2001. [7] R. Fletcher, A. Hakuli, and P. Imhof, Oil and Gas Journal, (2002) 54. [8] K.Y. Yung, R.J. Jonker, and B. Meijerink, ACS National Meeting August 2002 Boston, MA. [9] D. R. Rainer, E. Rautiainen, and P. Imhof, Applied Catalysis, submitted for publication. [ 10] E. Rautiainen and B. Nelissen, Hydrocarbon Engineering, submitted for publication. [ 11] W.S. Wieland and D. Chung, Hydrocarbon Engineering, March 2002. [ 12] P. O'Connor and A.P. Humphries, "Accessibility of functional sites in FCC" ACS preprints vol. 38, no. 3, p. 598, 1993. [13] L.A.Gerritsen, H.N.J.Wijngaards, J.Verwoert and P.O'Connor, "Cyclic deactivation: A novel technique to simulate the deactivation of FCC catalysts in commercial units", Catalysis Today, 11 (1991) p 61-72 [14] P. O'Connor and A.C. Pouwels, "Realistic commercial catalyst testing in the Laboratory." Proceedings of 8th International Symposium on Large Chemical Plants, Royal Flemish Society of Engineers, Antwerp, October 1992. [15] R. Pimenta, A. R. Quinones, and P. Imhof, Proc. Akzo Nobel Catalyst Symposium, Noordwijk, The Netherlands, June 1998. [ 16] D. Rainer, E. Brevoord, J. Gonzalez, and P. Imhof, Appl. Catal. A, submitted for publication. [ 17] G. Yaluris, W.-C. Cheng, M. Peters, and L. Hunt, Proc. NPRA Annual Meeting, New Orleans, LA, 18-20 March 2001. [ 18] D. R. Rainer, C. Vadovic, E. Rautiainen, and B. Nelissen, in preparation.

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 02004 Elsevier B.V. All rights reserved.

Reduction of additives

177

emissions from FCCU regenerators with

D. M. Stockwell and C. P. Kelkar

Engelhard Corporation, Iselin, NJ, USA Environmental emissions, particularly NOx emissions, have been a growing concern for refiners in recent years. NOx emissions from the regenerator are primarily derived from nitrogen in the coke. This work describes the development of active oxides as catalysts for reduction of NOx. Chemisorption of NO was used as a screening tool to measure and rank performance of metal oxides. The fresh capacity, and more importantly the hydrothermal stability of the chemisorption capacity, of the oxides were considered. Performance of these active oxides formulations were also evaluated during regeneration of spent FCC catalyst in a fluid bed. The resulting materials reduced FCCU NOx in full burn regeneration by as much as 30-70%. o

1. INTRODUCTION Since 1970, the Environmental Protection Agency (EPA) has tracked emissions of the six principal air pollutants termed criteria pollutants - carbon monoxide (CO), nitrogen oxides (NOx), sulfur oxides (SOx), volatile organic compounds (VOC), particulate matter (PM) and lead. Emissions of all of these pollutants have decreased significantly, except NOx that has increased approximately 10% over this period [ 1]. NOx has been identified as the primary cause for formation of ground level ozone (smog) formed when NOx reacts with VOC's in the presence of heat and sunlight. In 1997, as part of the revision to the Clean Air Act, the EPA issued a stricter ozone standard of 0.08 ppm averaged over an 8-hour period, compared with the older standard of 0.12 ppm averaged over a 1-hour period. According to the EPA, motor vehicles account for 49% of the NOx emissions [2], utilities contribute about 27%, and industrial and commercial factories account for about 19% of the emissions. While petroleum refining represents only 5% of the total emissions, these emissions are often concentrated in small areas. Many refineries are located in

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so called "non-aRainment areas" and can significantly contribute to local concentration of NOx and the concomitant ozone. The 1997 revision of the Clean Air Act includes a new source review (NSR) requirement, applicable when refineries make modifications that increase emissions. NSR standards require facilities to apply best available control technology (BACT) or lowest achievable emissions reductions (LAER) when grandfathered units are expanded in a way that increases emissions. As of this writing nine refining companies, covering over 30% of the U.S. refining capacity, have voluntarily signed consent decrees with the EPA. The consent decrees will result in criteria pollutant emission reduction of nearly 150,000 tons per year, including NOx emissions reduction of nearly 20,000 tons per year. The EPA also is negotiating settlement agreements with several other companies comprising an additional 30% of the domestic refining capacity. In fluid catalytic cracking (FCC), an inventory of microspherical catalyst is circulated between a riser-reactor and regenerator. Hydrocarbon feed contacts the clean regenerated catalyst in the riser reactor at 1000-1200~ High boiling hydrocarbon feeds crack to produce lower boiling components such as light cycle oil, gasoline, C3-C4 hydrocarbons etc. Cracking reactions also deposits coke on the catalyst. The cracked products are separated from the spent catalyst by steam stripping. The catalyst is then transferred to the regenerator where the coke is exothermically burned to regenerate the cracking catalyst. Gaseous products of regeneration are CO, CO2, H20, SOx, NOx, HCN, and NH3. This paper will deal with additives that will reduce NOx emissions from a full bum regenerator. 1.1 Chemistry of NO, formation NOx in the regenerator could hypothetically be formed by either of two mechanisms- thermal NOx produced from the reaction of molecular nitrogen with oxygen or fuel NOx produced from the oxidation of nitrogen containing coke species. Simple thermodynamic calculations have shown that thermal NOx is not a significant contributor to total NOx [3, 4]. To verify this, Dishman et al. [5] regenerated catalyst coked with isobutylene and compared this to catalyst coked with nitrogen-containing gas oil. At comparable regeneration conditions, NO was not detected with the isobutylene coked catalyst but was detected with gas oil coked catalyst. This experiment demonstrated that NO is not produced by thermal fixation of nitrogen. Zhao et al. [6] observed that in a circulating pilot plant, nitrogen-free feed did not produce NOx emissions even in the presence of platinum promoter. This showed that even the local temperatures around the platinum promoter are not high enough to favor catalytic NOx fixation from N2. The chemistry of how coke-bound nitrogen is converted to NOx in an FCC is very complex and not well understood. However, analogies can be

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drawn to several other processes that involve combustion of similar carbonaceous precursors. NOx formation during combustion of coal has been well-investigated [7]. The nitrogen containing products were seen to be dependent on the type of coal and combustion conditions. For most coals, HCN was the primary gas phase N-containing species during pyrolysis. NI-I3 appeared at the same time as H2, suggesting that NH3 is a secondary product of HCN reduction. It is thought that cleavage of heterocyclic forms of nitrogen compounds might be the source of HCN. Lower rank coals with more amino groups did produce s o m e Nit 3 as a primary product. Oxygen at low partial pressures actually increased HCN and decreased NH3 as it scavenges the H2 otherwise available for reduction. At higher oxygen partial pressures both HCN and NH3 are oxidized to NO [7]. Regeneration of spent hydroprocessing catalyst also involves oxidation of nitrogen containing precursors from the catalyst surface [8, 9] and similar products were observed. Interestingly, total yields of NO, HCN and NH3 accounted for only about one-third of the total nitrogen in coke. This indicated that major portion of the nitrogen in coke is converted either directly or indirectly to molecular nitrogen. Recent environmental concerns have spawned an increasing interest in understanding NOx chemistry in FCCU's. There are two basic families of regeneration processes, full and partial burn. In partial combustion mode, exactly as the term implies, coke is partially burned off the catalyst. This is achieved by using less than the stochiometric level of air. Typical exit gas composition in partial bum combustion is: 0.4 % O2, 15% CO2, 12% H20, 200 ppm SO2, 500 ppm NH3, 100 ppm HCN, 4% CO and balance nitrogen [10]. For both environmental and economic reasons, this CO rich flue gas is sent to a CO boiler. In the boiler the reduced nitrogen species can be oxidized to NO. In the full combustion mode, coke is more completely burned off the catalyst. This is achieved by using greater than stochiometric amount of air and also by using platinum-based promoters to oxidize carbon monoxide to carbon dioxide. Typical exit gas composition in full burn combustion is: 1.0 wt% O2, 16% CO2, 100 ppm CO, 12% H20, 400 ppm NO, 15 ppm N20 [ 11 ]. It should be noted that the above gas compositions are meant to convey ranges and will vary from unit to unit. The presence of precious metals and an oxidizing environment increases the oxidation of the reduced nitrogen species to NO and N20. At regenerator temperatures thermodynamics favors NO over NO2. Therefore although the term NOx is commonly used in referring to emissions, the primary constituent is NO. Recently Dishman et al. [5] and Zhao et al. [6] have carried out regeneration studies to study the evolution of gas phase nitrogen species during spent FCC catalyst regeneration. Temperature programmed oxidation (TPO) of spent catalyst in both studies show that carbon oxides evolve first and NOx

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evolution is observed only after most of the coke has been oxidized. Dishman et al. [5] analyzed the spent catalyst for C and N during the regeneration process. Results of these experiments showed that nitrogen is removed throughout the regeneration. Therefore, it was believed that molecular nitrogen is more likely a secondary product from the reduction of NO either with CO or coke [5]. In addition to the pathways involving nitrogen species there are several other competing reactions. Each of these competing reactions can affect the nitrogen pathway by either occupying a catalytic site or by affecting the concentration of another reactant. The most critical one is the CO promotion pathway. The presence of a platinum additive promotes oxidation of CO to CO2. A decrease in the CO concentration in the regenerator reduces the rate of the NO + CO reaction, causing the NOx emission from the regenerator to increase. NOx emissions could also increase because platinum is catalytically active for oxidation of NH3/HCN to NOx. Whichever mechanism applies the increase in NOx emissions when using a platinum-based CO combustion promoter has been observed in many commercial units.

1.2 Strategies for reducing NOx Several approaches are available to deal with NOx. These are (a) feed hydrotreating to remove NOx precursors, (b) FCC process hardware for reducing NOx formation, (c) catalytic approaches compatible with the FCC reactor which will either suppress NOx formation or catalyze its reduction, and (d) "end of the tailpipe" type solutions like SCR/SNCR. Feed hydrotreating is the most expensive approach. It also limits the refiner's flexibility in terms of being able to use different feedstocks and is probably not viable strictly from a NOx perspective. Process approaches focus on modifying the regenerator to facilitate better contact between NO and CO or coke. An example is Kellogg's Orthoflow regenerator that is claimed to produce 60-80% lower NO• than other designs [ 12]. These types of options are not immediately implementable by a refiner for an existing unit and may require drivers in addition to NOx reduction. Selective catalytic reduction (SCR) is a proven approach to reduce NOx with NH3. It has well-established history in reducing NOx in gas turbines and coal fired power plants but needs to be properly evaluated for FCCU' s. Catalytic approaches that can be employed inside an FCC are a very cost efficient and rapid way to reduce NOx. These materials are commonly referred to as additives in FCC literature, due to small amounts required to be blended in the circulating FCC inventory, but these additives are catalysts in their own right. A comprehensive review of all the additives for NOx reduction is not within the scope of this paper, but a review of different additives evaluated prior to 1998 can be found in Chen et al. [3]. A number of different classes of materials, such as perovskite- based, spinel-based, lanthanum or yttrium oxides,

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titanates, copper impregnated on ZSM-5, and ceria-based materials [6] have found to reduce NOx. Ceria based materials utilize the redox ability of ceria to cycle between C e 3§ and C e 4+ oxidation states. Ceria is reduced in the riser and serves as an agent to reduce NO in the regenerator. The presence of reducing agents in the regenerator will be required to maintain the ceria in the +3 state. Copper oxide is also part of many NOx reducing additive formulations [13]. Mixtures of NO and propane used to simulate reducing conditions were shown to reduce NO with high activity. As oxygen was introduced into the mixture the NO conversion activity declined. This could be due to the oxidation of hydrocarbon being the preferred route and the desired reductant no longer available for NOx reduction. 2. E X P E R I M E N T A L

2.1 Catalysts We prepared a series of additive formulations. An alumina support with particle and attrition characteristics suitable for use in cracking applications was used. Different active components were deposited on the support by incipient wetness impregnation of nitrate salts. A constant level of copper oxide was included in the formulations. After drying overnight, the catalysts were calcined at 1200 ~ F for 2 h. For additive D where the solution of total nitrate required exceeded the pore volume, the process was repeated with a calcination step in between. Deactivation of the additives was done by steaming at 1500~ for 4 h in 100% steam. Chemical and physical characterization of the different additives evaluated in this work is shown in Table 1.

2.2 Testing Catalytic performance of the additive was evaluated by NO chemisorption, NO/CO testing and temperature programmed oxidation of coked cracking catalyst. Table 1 Chemical and physical characteristics of different catalysts Catalyst

Composition

Surface Area (m2.~-1)

Pore Volume

A B C D

9.6 wt% CeO2 5.5 wt% Active oxide 10.5 wt% Active oxide 20.7 wt% Active oxide

90 75 92 78

0.42 0.39 0.31 0.32

(CC.8=1)

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2.2.1 NO chemisorption NO chemisorptions were carried out in a fixed bed reactor. The reactor was connected to a manifold that permitted us to feed NO containing gas stream to the reactor after appropriate pretreatment. Typically 1.0 g of additive was blended with inerts to make 2 g of total solids in the reactor. The catalyst was pre-reduced in H2 at 1300~ Upon cooling down to room temperature (< 80~ flow was switched to an argon stream containing 235 ppm NO. We used both a chemiluminescence meter and a quadropole mass spectrometer to measure the gaseous effluent during chemisorption. Both fresh and deactivated additives were tested for NO uptakes.

2.2.2 NO~CO testing Steady state testing of NO/CO chemistry was carried out in a fixed bed reactor at 1300~ for the deactivated additives. After pre-reduction in H2 at 1300~ and sweeping with Ar, flow was switched to a gas stream consisting of 235 ppm NO and 400 ppm CO at 1300~ This type of test at slightly "rich" condition has been described earlier [14]. At these conditions spent promoted equilibrium catalyst by itself yielded 0% NO conversion. Compare this to previous work [5] which showed that exposing spent equilibrium cracking catalyst; to a stream consisting of 124 ppm NO/ 10% CO gave ~100% NO conversion.

2.2.3 TPO of coked catalyst Coking and regeneration experiments were performed in a fixed fluid bed reactor. Coke was deposited with a light gas oil (CCR = 0.39%) containing 938 ppm total nitrogen at 970~ Regeneration was done in a 19% O2 containing gas stream while ramping from 970 ~ to 1300~ Carbon and nitrogen oxides generated during the TPO were measured by infrared and chemiluminescence cells respectively. Similar type of test with some variation has been reported earlier [ 15]. Each deactivated additive was blended to a 2 wt% level with a promoted commercial equilibrium catalyst (ECAT). The ECAT had a surface area of 171 mZ.g-1 with a zeolite surface area of 95 mZ.gl and contained 262 ppm and 615 ppm Ni and V respectively. 9 g of catalyst/additive blend was used for each experiment. Different amounts of coke were deposited on the catalyst by changing the amount of oil fed to the reactor.

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N2

= 2e-11 le-li Switch to N O / A r g o n

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0"10:00 0:15:00 Time, rain

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Fig. 1- NO chemisorption over bulk ceria 3. RESULTS & DISCUSSION 3.1 NO titration We hypothesize that NO reduction takes place via a Langmuir-Hinshelwood type of mechanism involving dissociative adsorption of NO. This adsorption tendency for NO was evaluated in the laboratory for many different materials. Typical results are shown in Figure 1. Introduction of NO gas causes the N2 signal to immediately and sharply increase. This shows that dissociative adsorption of NO and recombination and desorption of molecular nitrogen are facile, even at room temperature. As the surface of the material is saturated with the left over adsorbed oxygen, breakthrough of NO occurs. The areas over the curve for NO after subtracting for the appropriate physical lag measures the total number of sites available for dissociative adsorption. In the regenerator, CO will scavenge the adsorbed oxygen, producing CO2. This in turn makes the site available for further reaction, thereby completing the catalytic cycle. NO uptake capacity of commercially available CeO2 was measured (Table 2). The cerium oxide was then deactivated by calcining or steaming at 1500~

Table 2 NO titration of calcined and steamed cerium oxide NO uptake mmol. g-1

NO uptake retention (%)

Surface area retention

(%) As-is Calcined @ 1500~ Steamed @ 1500~

0.233 0.046 0.035

20 15

11 7

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These results are also shown in Table 2. There is an 80% loss of NO uptake capacity upon calcination and further additional 5% loss of capacity upon steaming. At the harsh conditions of steam and temperature prevailing in the regenerator, catalytic materials deactivate. Additives are no different. Hence additives with higher steamed NO uptake capacities, i.e. more sites, should have improved performance. At room temperature, the sites we are titrating are probably mainly the surface sites with very little sub-surface oxidation. At higher temperatures, NO uptakes would be higher due to sub-surface oxidation. We believe that in a regenerator, oxidation-reduction cycling of a site is rapid and only surface sites take part in NOx reduction. NO uptakes were measured on flesh and steamed additive A through D. Results are shown in Table 3. Note that A is a base formulation containing only cerium and copper oxides. Additives B through D are proprietary active oxide formulations with sequentially increasing loadings that also have a constant level of copper oxide. Surface area retentions in these cases are primarily a measure of surface area loss of the support. Note that additive A has slightly more than ten times the steamed NO uptake of bulk ceria on a per unit of ceria basis. This is expected if only surface sites participate, since supported ceria should have better dispersion. Additionally, copper oxide is also present in additive A, which will contribute some to the uptake. In this work, we have not tried to separate the effects copper oxide and ceria and their dispersions. Building on the deactivation hypothesis, proprietary active oxide formulations were developed with steam stability features superior to ceria. Additive B had flesh NO uptake capacity approximately half that of A. This is roughly in line with half the oxide content for B. The important difference is the steamed NO uptakes are almost identical. Therefore we see that these proprietary oxides can provide same number of sites for NO uptake at half the oxide loading. This could translate to a refiner needing to use less additive to achieve a constant NOx reduction. Table 3 NO uptake capacities of different steamed additives Catalyst

NO uptake flesh, mmol.g- 1

A B C D

0.045 0.018 0.038 0.078

NO uptake steamed, mmol.g-1 0.010 0.012 0.024 0.046

Uptake Retention (%) 22 65 62 59

Surface area retention (%) 84 85 76 74

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Doubling the active oxide loading from 5.5 (B) to 10.5 (C) doubled both the flesh and steamed NO uptakes. Further doubling the active oxide loading from 10.5 to 20.7, further doubled the uptakes. This suggests that the dispersion of active oxide is remaining fairly constant over a wide range of loadings, and that the relative dispersions are not being affected by deactivation. Note that additives A and C have very similar oxide loadings and show very similar fresh uptakes. Once again, the improvement is evident in the steamed uptakes. Additive C has twice the uptake as that of A. It is expected that providing sites for NO dissociation will lead to NOx reduction. If these sites deactivate rapidly, additive performance will also degrade. By providing active oxides with sites for NO reduction that are less prone to deactivation, the performance of the additive should be improved relative to ceria-based additive at constant loadings. 3.2 NO/CO testing Steamed additives were also tested for NO conversions under slightly fuelrich conditions, as described earlier. NO conversions at two different weights of additives in the blend are shown in Table 4. Considering that additive B had a slightly higher NO uptake than A, it is expected that B would display slightly higher NO conversions. Although the NO conversions with both the formulations are very close, additive A displayed somewhat higher NO conversion at both active weights. We believe the differences are small but we are continuing to investigate the reason tbr it. Systematically increasing the loading of active oxide from additive B through D increased NO conversion as would be expected. The similarity between results from NO uptake and NO/CO conversion are expected if only surface sites are involved. Both the tests are just different ways of measuring sites for NO dissociation. In one case we are measuring how many sites are available for dissociation of NO. In the second case the oxygen atom left over after the dissociative adsorption of NO is being rapidly scavenged by CO thus regenerating the site. 3.3 TPO studies Cracking catalyst regeneration studies were used to further test additive Table 4 % NO conversions at two different weights of additives for the NO/CO test Catalyst A B C D

0.05 g 16 12 23 36

0.15 g 62 56 72 85

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formulations. The test employs the full sequence of reactions of nitrogen deposition in coke via cracking gas oil, through NOx generation from coke during coke burning in air. It also employs realistic conditions for oxygen, from its compositional gradients due to the introduction of air near the bottom of a fluid bed, to the formation of oxygen-rich bubbles in the midst of carboncontaining catalyst, and the nearly complete consumption of oxygen during coke combustion. Although not shown here, the profiles of CO, CO2, NO and O2 look similar to the earlier work of Dishman et al. [5] and Zhao et al. [6]. Figure 2 shows total integrated NOx evolved vs. total integrated coke regenerated during a run. This data was generated from a conventional C/O study. As C/O was decreased from 9 to 3, total coke on catalyst increased from about 45 to 60 mg. As more coke is deposited on the catalyst, more NOx is emitted during the regeneration. Assuming that the carbon to nitrogen ratio of the coke is not changing dramatically with C/O, it makes intuitive sense that more NOx is emitted from catalyst with more coke. ECAT makes ~15 to ~20 gmoles of NOx with increasing coke in the C/O study. The same C/O experiment was repeated but with the ECAT blended with 2% by weight of deactivated additives, A through D. All the additives lowered NOx emissions compared to the ECAT. Additives A and B, both demonstrated 30-35% NOx reduction over ECAT. Increasing the weight of active oxide loading from 5.5 wt% to 10.5 wt% further decreased NOx by an additional 7 % for a total of 41% reduction over ECAT. Subsequent increases in loading of the active oxide to 21 wt% (D) did not further lower NOx emissions. Compared to the previous two tests, which relied on NO as a target molecule, this test involves regeneration of FCC coke. This in essence permits the in situ formation of NOx, as it would be in a commercial unit. It is pleasing to note some similarities between the different tests. Additive A and B have similar number of sites available for NO decomposition. The NOx reductions obtained in the TPO test were also similar. Doubling the active oxide loading from B to C to D had a proportional response on the number of sites. The NOx reduction due to additive C in the TPO test was also higher, but returns were clearly diminishing. Additive D, for example did not lead to any further NOx reduction. Apparently the higher number of active sites could not be fully utilized for some reason in the TPO test. On the other hand, in one commercial trial, we did test additives with two different oxide loadings and saw a more proportionate response.

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20 ECAT

15

c

O

B

A

i

II

i I

Z

10 J

i

40

i

r

r

i

t

45

50

55 Coke, mg

60

65

70

Fig. 2: TPO of coked blends of ECAT's with different catalysts Analysis of cracking yields done as part of the TPO experiments showed a small increase in H2 yields (+0.02 wt %) at constant conversion. There was no impact for increasing the loading of active oxides on hydrogen yields and the debit was constant across all the tested formulations. 3.4 C o m m e r c i a l trial

A refinery trial of the new NOx control additive was performed on a full bum unit that has an average baseline NO~ level of approximately 65 ppm and uses oxygen enrichment. During the trial period the exit oxygen concentration was maintained between 0.4 to 0.6 wt% and the total feed nitrogen was around 550 ppm. The unit continued to use a conventional platinum-based CO combustion promoter during the trial to control aiterbum at a relatively constant level. We observed about 48% reduction with a formulation similar to those described in this paper. A statistical process model confirmed these results [ 16]. The unit observed no noticeable change in coke or dry gas yields during the additive trial. 4. CONCLUSIONS We have developed new, proprietary additives for NOx reduction that have improved steam stability compared to previously known ceria materials. The improved steam stability translates to better NOx reduction in an FCCU. We used chemisorption of NO at room temperature to titrate for NO decomposition sites. Steady state testing with slightly rich NO-CO mixture at regenerator temperatures validated the effectiveness of these sites. TPO of spent cracking catalyst blends was also used to comprehensively simulate regeneration

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chemistries in the laboratory. The TPO test confirmed improved performance of additives with enhanced steam stability, and commercial trials have validated these laboratory results.

REFERENCES [ 1] National Air Quality- 2000 Status & Trends, U. S. Environmental Protection Agency, September 2001. [2] NOx- How Nitrogen Oxides Affect the Way We Live and Breathe, EPA, Office of Air Quality Planning and Standards, September 1998. [3] W.-C. Cheng, G. Kim, A. W. Peters, X. Zhao, K. Rajagopalan, M. S. Ziebarth and C.J. Pereira, Catal. Rev.-Sci. Eng., 40 (1998), 39. [4] A. W. Peters, X. Zhao and G. D. Weatherbee, NPRA AM-95-99 (1995). [5] K. L. Dishman, P. K. Doolin and L. D. Tullock, Ind. Eng. Chem. Res., 37(1998), 4631. [6] X. Zhao, A. W. Peters and G. W. Weatherbee, Ind. Eng. Chem. Res., 36(1997), 4535 [7] M. A. Wojtowics, J. R. Pels and J. A. Moulijn, Fuels Process. Tech., 34 (1993), 1. [8] E. Furimsky, A. Siukola and A. Turenne, Ind. Eng. Chem. Res., 35(1996), 4406. [9] P. Zeuthen, P. Blom and F. E. Massoth, AppI. Catal., 78(1991), 265. [ 10] G. L. Johnson, N. C. Samish and D. M. Altrichter, US Patent 4,744,962 (1988). [ 11] D. A. Cooper and A. Emanuelsson, Ener. & Fuels, 6(1992), 172. [12] R. B. Miller, T. E. Johnson, C. R. Santner, A. A. Avidan and J. H. Beech, NPRA AM-9648 (1996). [ 13] A. Corma, A. E. Palomares, F. Rey, F. Marquez, J. Catal., 170 (1997), 140. [14] A. Bhattacharya, M. J. Foral, W. J. Reagan, US Patent 5,750,020 (1998). [15] A. A. Chin, US Patent 5,021,146 (1991). [16] C. P. Kelkar, D. M. Stockwell, S. Winkler, S. Tauster, J. A. Sexton, G. A. Cantley, J. P. Wick, NPRA AM-02-56 (2002).

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 02004 Elsevier B.V. All rights reserved.

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Oxygen partial pressure effects on vanadium mobility and catalyst deactivation in a simulated FCCU regenerator G. Krishnaiah, L. V. Langan, J. A. Rudesill, and W-C. Cheng GRACE Davison, W.R. Grace & C o . - Conn., 7500 Grace Dr., Columbia, MD 21044 The rate of interparticle vanadium transfer in FCC catalysts as a function of the degree of oxidation of the steaming environment has been determined. The rate of interparticle vanadium transfer is low when the atmosphere in the steamer is reducing, but increases dramatically as the mole % excess oxygen becomes greater than zero. High rate of vanadium transfer is accompanied by a greater loss of surface area and catalytic activity. XPS measurements show that the oxidation state of vanadium increases from +3 when the steamer atmosphere is reducing to +5 when the % excess oxygen is greater than zero. Surface vanadium concentration follows a similar trend. In addition, data from a commercial resid FCCU operating in a partial bum were analyzed. The results of the analyses show that defining a partial burn operation with the flue gas CO/CO2 ratio is not adequate for characterizing the vanadium oxidation state. Carbon on regenerated catalyst is a better marker for the vanadium oxidation state in the regenerator.

1. I N T R O D U C T I O N The Fluidized Catalytic Cracking Unit (FCCU) remains the "heart" of the refinery, even more so with the recent demand for clean fuels. The demand for high-octane gasoline (albeit low sulfur) has remained unchanged with increased emphasis on alkylate feed. Narrow refining margins and increasing crude prices are limiting the refiners processing flexibility. Crude run reductions and increased bottom of the barrel (resid) processing in the FCCU are natural outcomes of these factors.

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Resid processing in the cat cracker brings with it additional challenges. The poorer feed quality and increased Conradson Carbon Residue (CCR) results in decreased yield selectivity and increased coke yields. The higher coke yields require heat rejection options such as catalyst coolers or partial burn regenerator operation. Resid feeds bring along increased nickel and vanadium contaminants (as porphyrins) that deposit quantitatively on the catalyst, negatively impacting catalyst activity and yield selectivity. The dehydrogenation activity of nickel and vanadium results in increased coke and dry gas yields. In addition, vanadium deactivates the catalyst by attacking the active sites and destroying the zeolite crystal structure. Furthermore, unlike nickel, vanadium tends to be very mobile, increasing its potential for deactivating the catalyst. Nickel passivating technologies, such as antimony and specialized matrix materials, have allowed the refiner to process higher nickel containing feeds while minimizing the nickel related dehydrogenation activity. It has long been accepted though not quantified that vanadium is most mobile when it exists in the +5 oxidation state in the regenerator [ 1, 2]. Therefore it is reasonable to state that an effective catalyst management practice that would minimize catalyst deactivation with vanadium would ensure that the bulk of the vanadium is either maintained in the +4 or lower oxidation state or bound in a form that is not mobile. It also follows that one would expect vanadium mobility to be minimized in a partial burn operation (over a full burn operation). Significant earlier work by several investigators has established that vanadium is mobile in laboratory catalyst deactivations that include 540 C calcination followed by 760 C steaming [3]. The mobile V species is reported to prefer a basic surface like layered magnesium silicate and to form heat stable vanadates such as beta-Mg2V207 and Mg3V208 identified by Laser Raman Spectroscopy (LRS) [4,5]. It is asserted that the mobile V species indicated is H2V207 formed from steam and vanadia [6]. Ortho vanadates were not seen by either LRS or 51V NMR; thus suggesting that H3VO4may not be the mobile species [7,8]. Since A1VO4 is not thermally stable, it will not likely provide vanadium passivation [9]. The exact nature of the mobile V species awaits actual in-situ detection, and that remains a difficult analytical challenge. To the best of our knowledge, a systematic study evaluating vanadium mobility on catalysts between the partial and full burn modes has not been conducted. The objective of this work is to measure the rate of vanadium transfer as a function of the level of excess oxygen in the steaming environment. We conducted experiments in a modified Cyclic Propylene Steam (CPS) deactivation unit. The CPS deactivation method was introduced to simulate catalyst deactivation by metals in a commercial FCC unit [ 10,11 ].

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2. E X P E R I M E N T A L M E T H O D S Properties of two commercially manufactured fluid cracking catalysts are shown in Table 1. The catalysts were selected based on the difference in their skeletal density so as to allow for easy sink/float separation. Catalyst D (Donor) was impregnated with 6000 ppm V (from vanadium naphthenate) and calcined at 1100~ In a typical steaming experiment, a 100 gram sample, consisting of a 3 to 1 blend of vanadium impregnated Catalyst D and Catalyst A (Acceptor), was placed in a quartz reactor and immersed in a sand bath, maintained at 1400~ The sample was fluidized and steamed for varying lengths of time (1, 2, 4 and 24 hours) in a mixture of steam, nitrogen, air, and 5% propylene in nitrogen. The atmosphere in the steam consists of 60 v% steam and 39 v% nitrogen; the remaining 1% is a mixture of oxygen and propylene. The flowrate of the air and 5% propylene streams was adjusted to give a wide range of propylene/oxygen molar ratios. Complete combustion of one mol of propylene to CO2 and H20 consumes 4.5 mols of 02. We define: % Excess 02 - (mols of 02 - 4.5 x mols C3H6)/total mols of gas The % excess 02 concentration was varied from a deficient-3.7% (highly reducing) to an excess 1.07% (oxidizing). In each case, the total flows in the bed were adjusted to maintain the superficial gas velocity below 0.1 ft/sec.

Table 1. Fresh catalyst properties Catalyst A

Catalyst D

Catalyst D (Post

Blend: 75/25

6000 ppm V)

(D/A)

Chemical Analysis: ,,

A1203 RE203 Na20 Ni V Physical Analysis" SA Zeolite Matrix

9Wt.% 9Wt.% 9Wt.% 9ppm "ppm 9M 2 / G M 9M 2 / G M 9M 2 / G M

49.85 3.07 0.38 27 48

33.47 0.97 0.31 kl+k3.

On the other hand, the LCO and gas fraction in cracking of extra heavy gas oil show a clear temperature dependency. At higher temperature more LCO is converted and more gas is formed. These effects are coupled to one another. Since no substantial coke formation takes place after 50 ms a cracking scheme for the extra heavy gas oil can be derived as shown in Fig. 10a. This reaction scheme is also applicable for the aromatic gas oil, since the cracking chemistry is the same. This has been demonstrated in [5]. Clearly, a given part Of the HCO fraction does not crack. The results with the spiked hydrowax demonstrate that this 'uncrackable' fraction consist of aromatics. Activation of aromatics can be expected to be very difficult as a result of the resonance stabilisation by the ~-electrons in the benzene ring. Moreover, even if the benzene ring would be activated it does not allow the formation of an attractive product spectrum upon ~3-scission. Cracking reactions of the multi-substituted polycyclic aromatics of the feed will be limited to dealkylation of the polycyclic cores, which subsequently accumulate in the HCO fraction. As the majority of sulfur in the feed has an aromatic character, it is interesting to see how these species behave under the applied conditions. The amount of sulfur in the product fractions is only dependent on the conversion, not on the temperature. Most of the sulfur from the HCO fraction ends up in the LCO fraction. This is in agreement with what is observed in the sulfur spectra. Predominantly LCO range sulfur species are formed by dealkylation of HCO range alkylbenzothiophenes.

230

Fig. 10. The cracking model for hydrocarbons and sulfur species for an extra heavy gas oil. (a) hydrocarbons, and (b) sulfur species. ( ~ ) = primary reaction, ( ~ ) = secondary reaction, and (.........:~ ) = incidental reaction. The benzothiophene cores accumulate in the LCO fraction, and due to their aromatic characteristics they are unable to react further. The dealkylation of alkyl-dibenzothiophenes goes accompanied with the formation of dibenzothiophenes, which lie in the HCO range. For gasoline sulfur some secondary reactions are observed. Liberation of thiophene molecules out of the large polycyclic feed species cause that these species evolve in the gasoline fraction. However, at higher conversion level some sulfur starts to disappear again. As the sum of the sulfur in the HCO, LCO, and gasoline fraction decreases this demonstrates that gaseous sulfur species are formed, i.e. some overcracking of aliphatic gasoline sulfur occurs. The observation that for gasoline sulfur overcracking takes place, whereas for hydrocarbons this is not the case can be explained from the relative amounts of both. Conversion of some gasoline range molecules will hardly be noticed in the large gasoline fraction, whereas for sulfur it will. The cracking scheme for sulfur species is shown in Fig. 10b. The combination of the hydrocarbon and sulfur cracking schemes explain the observed behaviour for the sulfur concentrations in the product fractions. Up to 50 % conversion the HCO hydrocarbons and sulfur species crack with the same rate, and the concentration of HCO sulfur remains unaffected. After 50 %

231 conversion the reactive sulfur species have already cracked, whereas the HCO hydrocarbons participate in cracking reactions. The result is that the HCO fraction gets more concentrated with sulfur. The opposite is the case for gasoline sulfur. Although, the gasoline sulfur concentration decreases, it is not a result of cracking of sulfur species. Where for sulfur cracking the dominant route is from HCO to LCO, the most dominant route for hydrocarbons is from HCO to gasoline. The accumulated sulfur species in the gasoline fraction are, therefore, diluted with cracked hydrocarbons from the HCO fraction. The increase of sulfur concentration in the LCO fraction is a result of the accumulation of the benzothiophenes. Where the hydrocarbons in the LCO range are fairly unreactive, the concentration increases. The temperature dependence is a result of the temperature dependence of hydrocarbon cracking reactions from LCO to gas and not of the cracking of sulfur containing molecules in the LCO fraction. In Fig. 11 a reaction scheme is presented that summarises the cracking behaviour of aromatic- and organic sulfur species as discussed in this section. This figure shows that under the applied conditions the aromatic species do not crack. Only the side-chains of these species will dealkylate from the stable cores. Note that only aliphatic sulfur species crack to hydrocarbons and hydrogen sulfide, and that the thiophene-, benzothiophene- and dibenzothiophene cores accumulate in the gasoline-, LCO- and HCO fraction, respectively.

5.

Conclusion

In conclusion, the aromatic backbone will not crack under realistic FCC conditions. Only the paraffinic fraction of the feed and the alkyl groups associated to the benzene ring in aromatic compounds are susceptible to cracking reactions. This is not surprising as the cracking of the benzene ring is low because of the resonance stabilising properties of the n-electrons. This stabilisation hampers the activation by a Br0nsted acid and the subsequent weakening of the bond at the 13-position. As a consequence only the side-chains of the (poly)aromatic feed will break, leaving the polycyclic cores in the HCO fraction, which account for the 'uncrackable' fraction.

References [1] [2] [3]

ACEA World-Wide fuel charter (2002), http://www.acea.be/ J. Corella and E. FrancEs, "Fluid Catalytic Cracking Ih concepts in catalysis design", Ed. M.L. Ocelli, ACS Washington D.C. (1991), 165-182 M.A. Den Hollander, M. Makkee and J.A. Moulijn, Appl. Catal. A. 187 (1999), 3-12

232

[4]

[5]

X. Dupain, L.J. Rogier, E.D. Gamas, M. Makkee and J.A. Moulijn, Appl. Catal. A. 238 (2003), 223-238 X. Dupain, E.D. Gamas, R. Madon, C.P. Kelkar, M. Makkee and J.A. Moulijn, Fuel (2003), accepted for publication

Fig. 11. Reaction scheme for the cracking of a FCC feed. Coke formation is not included.

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 9 Elsevier B.V. All rights reserved.

233

Evaluating factors that affect FCC stripper behaviour in a laboratory fluidised-bed reactor C.E. Snape ~, M.C. Diaz ~, C.L. Wallace b# and B.J. M c G h e e b•

aNottingham Fuel and Energy Centre, School of Chemical, Environmental and Mining Engineering (SCHEME), University of Nottingham, University Park, Nottingham NG7 2RD, United Kingdom bUniversity of Strathclyde, Department of Pure and Applied Chemistry, Thomas Graham Building, 295 Cathedral Street,~Glasgow G1 1XL, United Kingdom # Currem address: Engineering Physical Sciences Research Council, Polaris House, North Star Avenue, Swindon SN2 lET, United Kingdom *Current address: School of Biomedical and Biological Sciences, Caledonian University, City Campus, Cowcaddens Road, Glasgow G4 0BA, United Kingdom

An experimental protocol based on a laboratory fluidised-bed reactor has been developed to determine how the yield and composition of coke and the associated catalyst surface area vary as a function of stripper conditions in fluid catalytic cracking (FCC). In this study, it is demonstrated that coke from the fluidised-bed reactor prior to stripping and from an actual FCC riser unit are very similar compositionally, including the amoums of soft (chloroform-soluble) coke physically entrapped within the catalyst pore structure and only released after demineralisation that ultimately control the final coke yields. It has also been confirmed that hard coke from both actual feeds and a model compound, nnonene used to simulate the formation of catalytic coke, are all highly aromatic after short stripping periods. Further, transferring the deactivated catalyst from the fluidised-bed reactor after very short stripping times to other reactors has provided a means to ascertain precisely the influence of gas flow rate on stripping. Finally, the stripping behaviour of a selection of commercial FCC catalysts has been compared in the fluidised-bed reactor.

234

1. I N T R O D U C T I O N It is now generally accepted that, as well as being formed via the actual cracking reactions, coke also arises from thermal and metal-mediated (NiN) reactions, together with the entrained products which are symptomatic of incomplete stripping and can contribute to the overall level of coke [1]. The entrained products increase the hydrogen content of the coke and the additional air requirement gives rise to excessively high temperatures in the regenerator and additional steam, which in turn contribute significantly to the deactivation of FCC catalysts. The highly dynamic situation within a FCC unit is further complicated by the thermal reactions, which occur in the stripper section and can affect the yield and structure of the chloroform-insoluble coke (hard coke). Tests were conducted in a microactivity test (MAT) and a fluidised-bed reactor to develop an experimental protocol to determine how the yield and composition of coke and the associated catalyst surface area vary as a function of stripper conditions in FCC [2,3]. In both reactors, the use of rapid quenching allowed the relatively short stripping times encountered in FCC units to be simulated, independently of cracking. To develop the procedure, low sulphur vacuum gas oils (VGOs) with a low metal equilibrium catalyst (E-cat) were used for stripping periods of up to 20 minutes. Surface areas were determined before and after removal of the soluble (soft) coke with chloroform. Further, hard coke concentrates were prepared via demineralisation with hydrofluoric and hydrochloric acids for charaeterisation by solid-state 13C nuclear magnetic resonance (NMR). This approach was successfully demonstrated for FCC refinery catalysts [4,5] where the cokes were found to be highly aromatic in character (carbon aromaticities > 0.95), but differences in feedstock composition were still reflected in the structure of the cokes. Significant variations were evident in the structure of both hard and soil coke during stripping. Although solid-state 13C NMR indicated that the hard coke becomes more highly condensed with prolonged stripping, the surface area reduction by the hard coke remains fairly constant for stripping periods in excess of c a . 5-10 minutes and is small (10 m 2 g-l) in relation to the loss of surface area from the soil coke. It was found that a further pool of sot~ (chloroform-soluble) coke is physically entrapped within the catalyst pore structure and is only released after demineralisation. In fact, this second so~ coke fraction is much more highly aromatic than the first fraction and ultimately controls the final coke yield. For the combination of E-cat and VGOs investigated previously [2,3], typically about half of the final hard coke content of nearly 1% w/w catalyst is

235

derived from this second soft coke fraction by carbonisation. The structural information obtained has been used to formulate a model for the stripping process where the entrapped fraction of soluble coke undergoes cracking in competition with coke formation and evaporative removal from the catalyst. The first aim of this contribution is to demonstrate that the coke from the fluidised-bed reactor prior to stripping and from an actual FCC riser unit are similar compositionally, including the amounts of soil (chloroform-soluble) coke physically entrapped within the catalyst pore structure that ultimately control the final coke yields. Secondly, to demonstrate that transferring the deactivated catalyst from the fluidised-bed reactor after very short stripping times to other reactors provides a means to ascertain precisely the influence of gas flow rate on stripping. Thirdly, it is shown how the stripping behaviour of commercial FCC catalysts can be compared in the fluidised-bed reactor. Finally, this investigation provides a comparison of the cokes obtained from actual feeds and n-nonene, a model compound used to simulate catalytic coke formation, have also been compared at short stripping times. It should be emphasised that the aim of the fluidised-bed procedure is to simulate stripping phenomena as opposed to provide a combined simulation of cracking and stripping that might be obtained with other laboratory procedures, such as that developed by Yung et al. [6]. 2. EXPERIMENTAL 2.1. Riser samples and characterisation Approximately 37 g of catalyst from the top and bottom of an actual riser unit was used for the demineralisation procedure. Atier demineralisation, the second soft coke fraction was extracted with chloroform from the coke concentrate. Carbon contents were determined for the catalysts before and atter extraction of the first soft coke fraction (soft coke I) and for the coke concentrate before and after extraction of the second soft coke fraction (sott coke II). The hard coke concentrates obtained after demineralisation and subsequent chloroform extraction were analysed by solid-state 13C NMR as described previously using Bruker MSL100 and DSX200 instruments [4,5]. Normal and dipolar dephasing (non-protonated aromatic carbon) spectra were obtained using the quantitatively reliable single pulse excitation technique. Additional information on the coke concentrates was obtained from monitoring the small quantities of volatiles evolved in probe mass spectrometry (MS). A selection of cokes produced from n-nonene (see following section) were also characterised by solid state 13C NMR, 2.2. Fluidised-bed reactor experiments The fluidised-bed procedure for stripping (presented in Figures 1 and 2) has been described previously [2,3]. ATests were conducted on actual feeds using 70 g of

236

catalyst with 75-200 ~m particles being used to limit elutriation from the bed. The catalyst was placed in the bed and the reactor was allowed to heat up to 520~ under the flow of the fluidising nitrogen (1.5 dm 3 min-l). A further flow of 1.5 dm 3 min-~ of nitrogen was used to assist feeding 15 g of two low sulphur VGOs into the centre of the bed. These VGOs contained 1.3 wt% sulphur but had different hydrogen to carbon ratios (H/C = 1.60 c f 1.75) and different viscosities. The reason for using these different VGOs was to generate sufficient samples to perform the stripping tests in the fixed-bed and fluidised-bed reactors. The feeding period was ca. 40 s and zero stripping time was estimated from the superficial gas velocities through the bed. The characteristics of the low metals E-cat used for these tests and those with n-nonene are listed in Table 1. 20 mm

PROBE CONNECTION

|-.~

~.5mm

j,

V I ~ 10ram ~

GASEOUS/LIQUID PRODUCT

EXIT

CONNECTION TO REACTOR 470 rnna N,~

'~"

/

~v L 160mm

~-Feed

SINTER TO SUPPORT CATALYST

330 mn~

CONNECTION TO PROBE Reactor

Probe

T-Piece

Fig. 1. Construction of the fluidised-bed reactor. As previously, the fluidised-bed reactor was quenched by physically removing the reactor from the furnace and the samples were recovered to carry out the off-line stripping tests. To evaluate catalytic coke laydown using n-nonene as a model feed a variety of runs with varying feed to catalyst (F/C) ratios were studied. Two methods were

237

used to alter the F/C ratios, one where the catalyst charge and the time-on-stream (TOS) were kept constant but the feed injection rate was varied (runs denoted by the VF prefLx) and another where the TOS varied but the oil injection rate and the catalyst charge remained the same (rtms denoted by the VR prefix, Table 2). In all cases 50 g of the low metals E-cat was used (Table 1).

~ ~--

NITROGEN~

"~

([~ ~

FEEDER

"~'~'~-'-"---- PROBE

GAS METER

'~

~ ~ Z CARBOLITE FI

DRY ICE TR~

[

TYRE LINING

- ONE

Fig 2. The fluidised-bed reactor system Table 1 Characteristics of low metals E-cat used Parameter

Low Metal E-Cat

% Re203 % A1203

0.82 36.0 826 855 65

Nickel ppm Vanadium ppm MAT Yield Hydrogen Factor

2.0

Tests were also conducted in the fluidised-bed reactor with five commercial Eeats to ascertain any differences that might exist in the stripping behaviour of a representative cross-section of E-cats and to correlate the results with those from the proprietary Akzo accessibility index (A.A.I.) test [8]. The characteristics of

238

the E-cats are presented in Table 2. Compared to previously, the nitrogen flow rate was reduced to not much above the minimum fluidising velocity with only the - 3 8 lain material (just a few percem) being disgarded, as opposed t o - 7 5 l~m in the work thus far. A stripping time of 3 minutes was selected to act as a reasonable discriminator for the five E-cats investigated.

Table 2 Details of coke laydown experiments for n-nonene Experimem

Oil Injection Rate (ml/min)

Time-onStream

Catalyst Feed Weight (g) Weight (g)

(n~) VF1 VF2* VF3 VF4 VR1 VR2 VR3* VR4 * idemical runs

1.0 2.0 3.0 4.0 instant 2.0 2.0 2.0

10 l0 10 10 1 5 10 15

50.6 50.9 51.5 50.8 51.5 52.0 50.9 51.4

8.8 15.4 24.2 27.6 10.2 8.6 15.4 26.1

Catalyst to Feed Ratio 5.8 3.3 2.1 1.8 5.0 6.0 3.3 2.0

Table 3 Characteristics of the five commercial equilibrium catalysts. E-Cat number 1 2 3 4 5

A.A.I. 8.2 6.3 7.4 3.9 3.9

Ni contem (ppm) 366 476 201 2644 2079

V coment (ppm) 2805 2865 226 3944 1223

A1203 (wt%) 44.95 41.95 38.73 41.39 43.94

2.3 Off-line stripping experiments A fixed-bed reactor that was heated resistively and has been described elsewhere [7] was used to carry out stripping tests again at 520~ for 10 minutes, but with 1 g of the catalyst deactivated with the less viscous VGO (H/C = 1.60) and stripped for 10 seconds in the fluidised-bed. The nitrogen flow rates were 0.01, 0.05, 0.1 and 1.0 dm 3 min -~ and the pressure was kept close to 1 bar. The samples deactivated with the more viscous VGO (H/C = 1.75) described in the previous section were placed back in the fluidised-bed reactor and stripped under steam and nitrogen at 520~ for 10 minutes with and without the easily

239

extractable fraction of coke present, which enables the elucidation of the effect of soft coke I on the final coke yield. 3. R E S U L T S AND DISCUSSION 3.1. Comparison of riser coke and coke from fluidised-bed reactor prior to stripping Table 4 lists the carbon contents for the E-cats from the riser and the low metals E-cat deactivated with the less viscous VGO and stripped for l0 seconds in the fluidised-bed. Although differem feedstocks were used, Table 4 indicates that there is a general similarity in the overall distribution of easily extractable and physically entrapped sott coke (I and II, respectively) and hard coke. In particular, the absolute amounts of hard and physically entrapped soft coke are very similar. Given that after stripping, hard coke contents are generally of the order of 1.0-1.3 wt% as opposed to 0.5-0.7 wt% on a carbon basis, this result provides further evidence that it is the fate of the physically entrapped soil coke that controls the final hard coke yield. If this fraction carbonises completely, coke yields approaching c a . 1.5% would be obtained. Final hard coke yields of close to 1.0% mean that the quantity of the physically entrapped soil coke lost by a combination of cracking to lighter products and evaporation is similar to that which carbonises. Table 4 Coke distribution in riser and fluidised-bed samples

Total carbon (as received) Soft coke I (chloroform extractables) Soft coke II (extracted after demineralisation)" Hard coke

Percemage of carbon (wt%) Riser Fluidised-bed Top Bottom 3.0 2.5 2.8 1.6 0.7 1.4 0.8 1.1 0.9

0.6

0.7

0.5

aThe ratio of the carbon comem of the initial coke concentrate after demineralisation to that after chloroform extraction is used to derive the quantity of soft coke II.

Table 5 lists the aromaticities derived by solid-state 13C NMR and the atomic hydrogen to carbon ratios of the hard coke concentrates obtained after demineralisation of the catalyst framework and chloroform extraction of soft coke II. The fractions of non-protonated aromatic carbon of the total aromatic carbon were obtained by dipolar dephasing and these values were used to estimate the fractions of bridgehead aromatic carbon as described previously [4,5]. The hard coke from the bottom of the riser is less aromatic than that from the top of the

240

riser. The aromaticity of the hard coke from the fluidised bed reactor is higher than that from the top of the riser and this may reflect that the laboratory sample may not exactly correspond to zero stripping time. It should be noted that, after stripping, hard coke concentrates have much higher aromaticities, being in excess of 0.95 for both refinery and laboratory samples [2-5]. Although all three coke concentrates contain similar proportions of nonprotonated aromatic carbon of the total aromatic carbon (Table 5), the fact that larger corrections are needed for aliphatic substitutents in aromatic rings as the proportion of aliphatic carbon increases means that the estimated fractions of bridgehead aromatic carbon increase with increasing aromaticity (Table 5). Taking a peri-condensation model, the differences in estimated fractions of bridgehead aromatic carbon correspond going on average from 4-5 rings for the riser bottom sample to 6-7 rings for the fluidised bed sample. This compares with in excess of 10 rings for stripped catalysts [4,5]. Table 4 Summary of structural parameters for the chloroform-insolublecoke concentrates Riser Atomic H/C ratio Aromaticity (+ 0.02) Fraction of non-prot, aromatic C (+ 0.03) Fraction of bridgehead aromatic C (+ 0.03)

Fluidised-bed

Top 0.73

Bottom 0.88

0.63

0.86 0.55 0.45

0.78 0.55 0.40

0.91 0.55 0.48

Figure 3 compares the probe MS obtained at the same temperature for the coke concentrates from the riser samples. The spectrum for the riser top sample displays peaks extending to m/z of 500 with a maximum close to 400. In contract, the riser bottom sample displays much weaker peaks in this mass range with ions having m/z ratios less than 120 arising from alkyl chains (e.g. 69 and 85) and phenyl moieties (e.g. 91 and 105). Since ions arising from PAHs in the m/z range 300-500 can be assigned to 6-10 ring systems, probe MS provides qualitative evidence to support the findings from solid state ~3C NMR that the sample from the top of the rise is more highly condensed. 3.2. n-Nonene cokes Figure 4 shows CP/MAS 13C NMR spectra of hard coke concentrates from nnonene coke after (a) zero time and (b) l0 minutes stripping. Table 5 compare the aromaticities for all the hard cokes prepared from n-nonene.

241

Riser Bottom

:l~ilT+

,.

II

I!

lm N.

i, .'~.

Riser Top

U. ~,

S.'

]I

169

II

'

'

II

Fig. 3 Probe MS for riser hard coke concentrates. (b)

Ca)

I

+S ~oli

+-f +5+

i+i PPl~

~

9

. . . . . . +.o . . . .

" ...... ,.+

'~',..

+o

= + +

. . . . .

Fig. 4 CP/MAS 1 3 C NMR spectra of hard coke concentrates from n-nonene coke atter (a) zero time (VR1) and (b) 10 minutes stripping.

242

Table 5 Summary of carbon aromaticities of coke concentrates from n-nonene experiments. Experiment VF1 VF2 VF3 VF4 VR1 VR2 VR3

Aromaticity (CP/MAS) 0.89 0.93 0.93 0.91 0.86 0.83 0.93

Aromaticity (SPE) 0.82 0.89 0.90 0.83

All the hard coke concentrates have high aromaticities and there is nothing to differentiate the one obtained at zero stripping time to all the others obtained after prolonged stripping periods. This trend mirrors exactly that reported previously where aromaticities for a VGO feed ranged from 0.86-0.90. This latest finding further substantives the hypothesis, that once formed, the initial hard coke cannot be cracked readily during stripping to yield substantially reduced quantities of hard coke.

3.3. Off-line stripping experiments Table 6 lists the stripping results obtained by placing the low metals E-cat coked with the more viscous VGO back in the fluidised-bed reactor. The initial carbon content of the catalyst prior stripping was 4.6 wt% and all stripping tests were performed at 520~ for 10 minutes. The high carbon content of the catalyst prior to stripping in comparison to that obtained previously (i.e. 4.6 cf. 2.8 wt%) might be attributed to the different characteristics of the feed and to the removal of the deactivated catalyst at zero stripping time in this case. Table 6 Stripping results obtained using the fluidised-bed reactor with steam and nitrogen Strippingagent

Flow rate (dm3 min-~)

Gas velocity (cm s-~)

Carboncontent~ (wt%)

Nitrogen Steam Steam

3.0 2.8 1.9

2.5 2.4 1.6

1.3 1.2 1.1

Nitrogen 3.0 Steam 2.9 ~Carbon contents determined with + 0.3% error.

2.5 2.5

1.1 0.8

Coked E-cat Before extraction After extraction

Steam causes a more significant reduction in carbon content than nitrogen, independently of the presence of soft coke I. However, the change in gas

243

velocity for the stripped E-cat with steam does not seem to alter the amount of remaining coke in the catalyst (ca. 1.2 wt%). On the other hand, the removal of soft coke I improves the efficiency (ca. 6% more) of the stripping process both in nitrogen and steam environments, suggesting that part of soft coke I contributes to the formation of hard coke and/or soft coke II. Clearly, the relative contributions from evaporative and cracking processes to the removal of the physically-entrapped soft coke have still to be evaluated, but the presence of steam, as opposed to nitrogen, could have a chemical influence based on the extensive literature on the pyrolysis of coals and oil shales [9]. The effect of nitrogen flow rate on the carbon contem obtained for the stripping experiments in the resistively fixed-bed reactor is presemed in Table 7. It can be seen that the carbon content decreases with increasing flow rate, reaching a constant value within experimemal error at flow rates higher than 0.1 dm 3 min-~. Comparing the values in Table 5 with that of the 10 minutes stripped catalyst in the fluidised-bed reactor (ca. 1.0 wt% carbon using a gas velocity of 2.5 cm s-l), approximately the same amount of carbon remains in the catalyst when using similar gas velocities. These results help to establish the sweep gas velocity required to achieve the maximum extem of stripping by removing imermediate cracking products from soft coke II and thus preveming their carbonisation. Table 7 Effect o f nitrogen flow rate on stripping in a fixed-bed reactor Nitrogen flow rate Gas velocity (dm 3 min -l) (cm S-l) 0.01 0.4 0.05 2.2 0.1 4.4 1 43.8 #Carbon comems determined with +0.3% error.

Carbon contem ~

(wt%) 1.7 1.1 1.0 1.0

3.3. Comparison of E-eats The residual carbon coments for the E-cats atter the 3 rain. stripping test in the fluidised-bed are listed in Table 8, which also gives the BET surface areas for the fresh and stripped catalysts. The surface areas are broadly consistent with the carbon coments with E-cat 1 giving the greatest recovery, E-cat 2 the least and E-cats 3, 4 and 5 being intermediate. In terms of ease of stripping, E-cat 1 is the best performing catalyst and E-cat 2 the worst. E-cat 1 also has the highest accessibility value in the A.A.I. test (Table 1). The relatively poor performance of E-cat 2 could possibly

244

be connected with a combination of poor transport and cracking either in the matrix mesopores or in the zeolite micropores. However, the significant differences in the catalyst compositions do not appear to have major influence on the carbon content prior to stripping. In particular, metals contents do not appear to bear any relation to coke yield for the 3 min. stripping tests. Table 8 Carbon contents and surface areas (SA) for 3 min. stripped E-cats in fluidised-bed reactor E-Cat number

Carbon content (wt%)

SA (m2 g-l) % SA As received Stripped recovery 1 1.33 ~ 115 93 81 2 3.20* 95 62 65 3 2.16 122 87 71 4 1.84 110 75 68 5 1.78 118 87 74 #Duplicate determinations on different samples of the same batch were 1.40 and 1.25%. *Mean of 6 determinations from the two duplicate runs. 4. C O N C L U S I O N S 1. Coke from the fluidised-bed reactor prior to stripping and from an actual FCC riser unit are very similar compositionally, including the amounts of soft (chloroform-soluble) coke physically entrapped within the catalyst pore structure that ultimately control the f'mal coke yields. 2. Both hard cokes from actual feeds and model compounds obtained at short stripping times are highly aromatic in character. Thus, once formed, it is unlikely that hard coke yields can be reduced via cracking as stripping progresses. 3. Transferring the deactivated catalyst from the fluidised bed reactor after very short stripping times to a fixed-bed reactor has provided a means to ascertain the influence of gas flow rate on stripping. Steam has been found to be a slightly more efficient stripping agent than nitrogen. In both reactors investigated, a critical gas velocity needs to be exceeded to minimise the final coke yield. 4. The ease of stripping behaviour in the fluidised-bed test can vary significantly for commercial E-cats. 5. A C K N O W L E D G E M E N T S The authors thank the Engineering & Physical Sciences Research Council (EPSRC) for financial support, and Shell, Akzo Nobel and BP Amoco for supplying the equilibrium catalysts used.

245

REFERENCES [1] P. O'Connor and A.C. Pouwels in Catalyst Deactivation 1994, B. Delmon and G.F. Froment (eds.), Elsevier, 1994 and Studies in Surface Science and Catalysis Vol. 88, Elsevier, 1994, 129 and references therein. [2] C.E. Snape, Y.R. Tyagi, M. Castro Diaz, S.C. Martin, P.J. Hall, R. Hughes and C.L. Koon, in Fluid Catalytic Cracking V. Materials and Technology Innovations, M.L. Occelli and P. O'Connor (eds.), Studies in Surface Science and Catalysis Vol. 134, Elsevier, 2001, 239. [3] C.E. Snape, Y.R. Tyagi, M. Castro Diaz, S.C. Martin, P.J. Hall, R. Hughes and C.L. Koon, Trans. I. Chem. E., No. 78A (2000) 738. [4] C.E. Snape, B.J. McGhee, J.M. Andresen, R. Hughes, C.L. Koon and G. Hutchings, Appl. Catal. A: General, 129 (1995) 125. [5] C.E. Snape, B.J. McGhee, S.C. Martin and J.M. Andresen, Catalysis Today, 37 (1997) 285. [6] K.Y. Yung, R.J. Jonker and B. Meijerink, Am. Chem. Soc. Div. of Pet. Chem., No. 47(3) (2002) 281. [7] G.D. Love, C.E. Snape, A.D. Carr and R.C. Houghton, Organic Geochemistry, 23 (1995) 981. [8] K.Y. Yung, P. Imhofand M. Baas, Prepr. Am. Chem. Soc. Div. of Pet. Chem., No. 47(3) (2002) 270. [9] E. Ekinci, A.E. Putun, M. Citiroglu, G.D. Love, C.J. Lafl~rty and C.E. Snape, Fuel, 71 (1992) 1511 and references therein.

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 02004 Elsevier B.V. All rights reserved.

247

Effect of FCC Variables on the Formation of Gasoline Gum Precursors William Richard Gilbert

PETROBRAS R&D Center (CENPES) Cidade Universit~ria, Quadra 7, 21949-900 Rio de Janeiro, Brazil. Tests done in a pilot riser unit showed how most of dienes that end up in the FCC gasoline are produced by post-riser thermal cracking reactions. The effect of feed quality was also investigated, showing that feed properties that define the potential conversion of a given feed also determine the stability of the gasoline produced. Adaptations to the experimental protocol of a MAT scale reactor are described that allow its use as a gasoline stability evaluation tool. 1. INTRODUCTION Gum in motor gasolines is produced from the reaction of unstable species present in the gasoline with oxygen in the air during storage. The reaction scheme is complex and several pathways may be involved, but in general it is agreed that the rate limiting step is the initiation of a chain of free-radical reactions [1 ]. Once gum is formed, damage to the motor is produced in places where the fuel evaporates, leading to deposits that eventually coke, such as injection systems and valves. Of the major components of the gasoline blend, the FCC is the greatest contributor to gum formation. With the tendency to process heavier feedstocks at higher severities in the FCC, the stability of the gasoline produced gets even worse. This is particularly worrying in countries like Brazil, where more than 70% of the gasoline pool is provided by the FCC. The tightening of gasoline specifications, specifically in the case of sulfur content, is driving many refiners to adopt hydrotreating of the FCC gasoline as the only way to attain the new levels. The

248

necessity to control gasoline stability is still valid in this scenario, as dienes in the FCC product lead to coking of the hydrotreater catalyst beds. The best way of testing gasoline stability is the measurement of the gum produced after three months dark storage at 40~ Since this method is time consuming and not practical for routine stability assessment, several more convenient methods have been developed, which either try to accelerate the gum formation process by exposing the gasoline to pure oxygen at high temperatures (100~ e.g. potential gmn (ASTM D 873) and induction period (ASTM D 525) or try to detect reaction intermediates or precursors such peroxides (ASTM D 3703) and dienes (UOP 326). Although these alternative methods may be adequate for refinery samples, new techniques had to be developed for small scale FCC reactors as will be described below. The standard solution for the gasoline stability problem has been to inject gum inhibitors and dilute the FCC gasoline with other more stable components in the blending process. These gum inhibitors often are free radical scavengers, such as hindered phenols or amines, and require careful control of mercaptans and HzS levels in the FCC gasoline post-treatment units to ensure their effectiveness. Although it has been known that things like feed quality and thermal cracking affect FCC gasoline stability, there is relatively little detail in the literature on how the FCC variables actually produce the observed results. For instance, the FCC operating experience shows that higher stability gasolines are produced by increasing catalyst/oil ratios (C/O) while keeping other variables constant, but when Riser temperature is raised, a change that also increases the C/O, stability deteriorates. In FCC there is a complex network of simultaneous reactions taking place, including both catalytic and thermal reactions. The balance and the timing between the two types of reactions is very important in determining the outcome of the gasoline composition and stability. A series of pilot riser experiments described in this work shows how thermal cracking taking place at the beginning of the reaction is not so important when compared to post-riser thermal cracking. Another set of results, also from the pilot riser, show that feed chemical composition rather than boiling point range or carbon residue is critical for gasoline stability. Lastly, adaptations to the experimental protocol of a small scale fluidized bed MAT type reactor (ACE unit) are described that allow its use as a gasoline stability evaluation tool. In a previous study [2], it was shown that the DCR pilot riser [3,4] was capable of reproducing the potential gum profile of PETROBRAS commercial FCCs, when run with the same catalyst, feedstock, riser and regenerator temperatures, therefore validating it for gasoline stability experiments. In the same

249

study it was also shown that gasoline potential gum could be correlated with diene concentration in the total liquid product, making TBP column fractionation of the pilot riser gasoline unnecessary for stability evaluation. 2. E X P E R I M E N T A L

In the pilot riser tests, a DCR [3,4] circulating unit was used which simulates commercial unit operation by using catalyst circulation to control an adiabatic riser exit temperature. Coke yield was calculated from the regenerator flue gas composition. Chromatographic Simulated Distillation of the syncrude, using 221 ~ and 344~ as cut points was used for the gasoline and LCO yield calculation. At least three runs at different C/O ratios (varying feed temperature) at each of the conditions was done to establish a curve from which results at iso-conversion could be interpolated. Conversion in all experiments was defined as the added yields of coke, dry gas, LPG and gasoline. To determine the gasoline stability in the first runs, analysis of the diene content directly on the syn-crude was done using UOP 326 procedure. When diene determination by supercritical fluid chromatography (SFC) became available, it was used instead. Table 1 summarizes the test conditions. Riser temperature refers to the internal temperature in the riser exit. Stripper and disengager vessel temperatures are wall temperatures controlled by electric heating. In this series, catalyst A and VGO2 were used. The disengager vessel receives the hydrocarbons from both the stripper and the Riser before sending them to product recovery, and drains the catalyst from the riser into the stripper. Table 2 and 3 show selected properties of equilibrium catalysts and feeds used in the runs. In Table 3 AtR stands for atmospheric resid, KVGO for a 50% mixture of coker gasoil in vacuum gasoil, DO for a 50% mixture of decanted oil in vacuum gasoil. In the second series of pilot riser tests, catalyst A was run with the different feeds in Table 2, at constant riser temperature (540~ to investigate the feed quality effect on gasoline stability.

250

Table 1 Temperature dependence of Gasoline Stability Study Run Riser T Stripp. T Diseng.T Comments 1 540~ 540~ 540~ Base case (low temperature) Overall high temperature case 2 550~ 550~ 550~ Low disengaging vessel temp. 3 550~ 550~ 540~ Low stripper temperature 4 550~ 540~ 550~ Low post-riser temperature 5 550~ 540~ 540~ Table 2 Catalysts properties Catalyst A A1203 %p 31.1 RE203 %p 2.04 Ni ppm 1825 V ppm 1523 SA m2/g 188 MSA m2/g 42 MiPVol cm3/g 0.068 MAT %p 73

B 36.9 2.2 3365 2555 121 34 0.041 58

In all tables and graphs, conversion is defined as the sum of the yields o f dry gas, LPG, gasoline and coke, and second order conversion as the ratio o f the conversion divided by 100% minus conversion. Table 3 Properties of feedstocks used in the study atrl atr2 atr3 atr4 atr5 atr6 kvgo Feed 13.2 18.6 27.9 2 8 . 1 1 8 . 2 1 1 . 9 16.9 API gravity 92.4 104.6 109.8 115.6 87.5 83.4 69.7 Aniline Pt ~ 2044 1068 9 4 . 3 646 1 1 3 6 407 2196 Basic Nitr. ppm 573 567 501 510 475 577 439 T50 SimDist ~ 9.0 8.0 1.9 3.8 3.4 1 3 . 3 0.54 RCR w% * NMR Sat.nyd.% t 95.2 96.4 9 7 . 3 97.6 95.4 95.4 9 0 . 3 * Carbon residue measured by method ASTM D 524. t NMR Sat.Hyd.% is the saturated hydrogen % determined by NMR.

do 9.8 58.3 968 433 3.8 86.0

vgol 19.5 80.6 486 458 0.9 94.0

vgo2 19.1 76.1 1047 443 0.37 96.1

vgo3 18.5 88.2 1019 499 1.84 95.1

In addition to the pilot riser tests, two series o f tests were done in the A C E unit [5], varying reactor temperature between 540~ and 560~ and analyzing diene

251

concentration in the total liquid product using SFC chromatography. In the first runs, the standard test protocol was used, where product exit line was kept constant at 530~ In the second series, the product exit line was purposefully maintained at the same level as the reactor temperature. Feed VGO3 and catalyst B were used in both cases. SFC was done in a Bergers Instruments chromatograph using CO2 at 244bar and 40~ as the carrier fluid in a silica column with a Agilent 1100 UV detector. Diene concentration in the gasoline was calculated by dividing the value measured in the total liquid product by the weight percent of gasoline in the liquid.

3. RESULTS AND DISCUSSION Raising riser temperature produced the expected decrease in coke yield and increase in dry gas yield (Table 4). The last result was interpreted as a sign of more thermal cracking in the riser. Aromatic content of the gasoline also increased with riser temperature. Gum formation reactions are controlled by their initiation rates, which in turn depend on the presence of precursors such as dienes, therefore, small increases in diene concentration have a dramatic effect on gum formation. When the diene concentration in gasoline is examined (Table 4 and figure 1), it shows that despite the higher dry gas yields in all the runs where riser temperature was 10~ above the base case, there was an increase in diene concentration only in the runs where the disengager vessel temperature was also high (run #2 and #4). To explain this it is assumed that the conjugated dienes produced from thermal cracking, being very reactive, are rapidly consumed in the catalyst mediated reactions. When no catalyst is available, as is the case of the disengager vessel, any diene that is produced ends up in the product. Figure 1 also shows the effect of C/O mediated conversion on the gasoline stability, conf'mning the rule of thumb that higher C/O produce more stable gasolines.

252

Table 4 Yield profiles and gasoline quality results from the pilot riser temperature dependence tests. Run 1 2 3 4 5 Riser T ~ 540 550 550 550 550 Stripper T ~ 540 550 550 540 540 Disengager T ~ 540 550 540 550 540 Conversion w% 75 75 75 75 75 Coke w% 6.6 6.0 6.3 5.8 6.3 Dry Gas w% 3.8 4.7 4.3 4.1 4.1 LPG w~ 18.6 19.1 17.7 19.5 18.0 Gasoline W% 47.6 45.9 46.6 45.6 46.7 LCO w% 14.3 14.2 14.2 13.8 14.3 Bottoms w% 10.7 10.8 10.8 11.2 10.7 GLN Aromatics w% * 34.1 36.7 35.0 34.8 35.6 GLN Olefins w% 33.8 33.3 35.0 34.7 34.7 GLN Saturates w% 32.2 30.0 30.1 28.0 27.7 Dienes 12 mg/100ml 6.91 9.28 7.24 8.91 7.19 * Gasoline composition measured by PIANO GC analysis, diene concentration measured by the UOP 326 method. 10.0 .................................................................................................... o A

o

)1(

e~

r

7.5

t5 .0

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

70

75 Conversion wt%

80

Fig. 1. Pilot Riser gasoline diene concentration as a function of conversion and temperature. (11) Base Case, (A) High Disengager vessel temperature, (:g) High Stripper temperature, (O) All temperatures high, (+) Low post-Riser temperatures.

In the second set of pilot riser experiments feed quality and C/O were varied at constant temperature (540~ differences in conversion and coke yield were so great between the feeds that an iso-conversion table for the comparison of selectivities was not possible. Nevertheless the results clearly showed that the

253

potemial conversion, translated in the correlation (equation 1) as the percentage of saturated hydrogen H s a t (measured by NMR) and the basic nitrogen content N b a s i c , was the determining factor of diene concentration and whence of the gasoline stability. Feeds with either high boiling range (AtR2 in Table 2) or high carbon residue (AtR6) produced good to moderate diene concentrations, lighter but more aromatic and contaminated feeds on the other hand (OD and KVGO) were the worse performers both in conversion and gasoline stability. It is interesting to notice that C/O ratio which is one of the most statistically important factors in conversion (Eq. 2), only appears in the diene correlation in combination with feed quality, as the product of basic nitrogen times C/O. In other words, as the concentration of basic nitrogen decreases in the feed, the importance of the C/O in the gasoline diene concentration becomes less perceptible. Dienes

= 95 - 0.96-

HSat +

Conversion = 1.7. C /0 -

0.0041-

0.003.

Nbasic -

Nbasic +

0.00034.

2.4.

HSat

Nbasic . C/0

- 170

(1) (2)

The measurement of diene concentration in the total liquid product of the FCC pilot riser requires a 50ml sample, which would preclude the use of MAT scale reactors in gasoline stability studies. With the development of the SFC method for diene determination (figure 2), this constraint was lifted and a series of experiments was done varying the ACE unit operating conditions until it could be shown that the gasoline stability produced would respond in the same way to changes in reactor temperature as in the pilot riser. The ACE unit (figure 3) is a fluidized bed reactor placed inside a temperature controlled oven, where feed is injected at the bottom and the products are collected at the top through a side line, which in the standard protocol is kept at 530~ regardless of the reactor temperature. In the first experiment reactor temperature was varied between 540~ and 560~ but the product line was kept at 530~ No increase in diene concentration could be detected (figure 4). Figure 4 also shows that the diene concentration values for the ACE unit were well below those of the pilot riser, for the same feed-catalyst pair. Suspecting that the lower temperature in the product recovery line could be acting like a quench, in the same way that happened in the cooling of the pilot riser disengager vessel in the experiments described before, the product line temperature control was set at the same value as the reactor temperature and a new series of tests varying reactor temperature was performed. This time diene concentration responded in the expected way (figure 5), with dienes concentration increasing as the reactor temperature is raised. As residence

254

time is k n o w n to be an important factor in thermal cracking reactions, the effect o f the position o f the product collector was also verified in the second series of tests.

~D r ~-,

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.

.

.

.

.

.

.

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.

.

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TE135

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255

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Conversion/(100-Conversion) Fig. 4. Diene concentration in the ACE unit gasoli-ne using the standard test protocol. (O) ACE 540~ (121)ACE 550~ (A) ACE 560~ (X) Pilot Riser 540~ (g) Pilot Riser 550~

O

0 " 7 1J~- ~

?5

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ii

0.5 0.3

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With the new procedure the results in the ACE unit were closer to those of the pilot riser. The two effects, C/O and reactor temperature can be shown in figure 4. The difference in diene concentration from 540~ to 550~ appears to be greater

256

then 550~ to 560~ This could be because of the difficulty in maintaining the whole of the product line at 560~ as other thermocouples in the line, particularly at the end of the line were cooler then the control thermocouple, and the difference increased with the temperature. The influence of the collector position wasn't statistically significant, showing that differences in product residence time in the product line were very small. 4. CONCLUSIONS The pilot riser tests, later confirmed by the ACE unit results, show that thermal cracking happening after the main catalytic reactions and catalyst stripping is critical for gasoline stability. Gum precursors such as dienes may be formed from thermal cracking occurring at any time in the process, but their inherent reactivity means that most of them are consumed in the FCC catalytic reactions, so long as there is catalyst present. Feed quality is one of the most important variables affecting gasoline stability. The same feed properties that def'lne the potential of conversion for a given feed, e.g. saturated hydrogen and contaminant basic nitrogen, are also determining factors in gasoline stability. The worse the feed quality, the more relevant other operating variables, such as C/O ratio, become. This is particularly evident when refiners try to compensate for the low conversion of refractory feeds by increasing reaction severity (riser temperature), which aggravates the gasoline stability problem. With small changes to the test protocol and the use of SFC for diene detection in the FCC total liquid product, it is possible to use a MAT scale laboratory test reactor, such as the ACE unit, in the investigation of gasoline stability problems.

REFERENCES

[1]. M. W. Scherpfer, C. A. Stansky, FL-81-79, NPRA Houston, November 1981. [2]. www.gracedavison.com, 2001. [3]. www.zeton,com, 2001. [4]. W. R. Gilbert,, TT125, 11th Brazilian Catalysis Conference, Bento Goncalves, 2001. [5]. www.kaysertech.com, 2001.

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 02004 Elsevier B.V. All rights reserved.

257

Distributed Matrix Structures- novel technology for high performance in short contact time FCC D. M. Stockwell, X. Liu, P. Nagel, P. J. Nelson, T. A. Gegan, and C. F. Keweshan

Engelhard Corporation, 101 Wood Avenue, Iselin, NJ 08830 Engelhard has introduced a new FCC technology platform called DMS (Distributed Matrix Structures), which features several catalyst and additive products. In this paper we report on the technology features behind these catalysts. Model matrix materials were synthesized and porosimetry results showed that a continuum of DMS mesopore diameters and surface areas could be obtained. Diffraction and SEM data confirmed this was due to the presence of 14-50 nm DMS crystallites of a synthesized silica-alumina. Silica-stabilized 7-A1203 was also present when the DMS crystals were small. These materials contain mainly Lewis acid sites, a nearly constant acid site density of about 5 gmol m -z, a n d significant cracking activity. Macropore volume and matrix dispersion were improved by employing ultrafine particle size precursors and a house-of-cards morphology. When synthesized with zeolite-on-matrix morphology, these DMS materials have led to dramatically improved yields in short contact time (SCT) FCCs. 1

INTRODUCTION

Engelhard's Distributed Matrix Structures (DMS) technology platform features several catalyst and additive products for specific applications. NaphthaMax | the first product from this platform, was developed expressly for modem FCC designs featuring short contact time cracking [1, 2]. Other products which have been developed include NaphthaMax | -LSG for gasoline sulfur reduction [3], Endurance TM and Flex-Tec T M for resid feeds, and Converter TM - the first conversion additive for FCC [4]. As of mid-2003, products employing DMS technology were used in over 70 commercial FCC

258

units, and have given impressive performance improvements in these applications. DMS catalyst systems employ a novel macroporous zeolite-on-matrix morphology [1], which has consistently given rise to 15-25% lower coke selectivity in commercial gas oil cracking than competitive technologies [2]. Catalytic coke is reduced by the pre-cracking of the feed on the highly selective external surfaces of the exposed zeolite. This structure minimizes diffusion path length from the pre-cracking event to the internal zeolitic cracking sites. The macroporosity aspect is also particularly important for Short Contact Time (SCT) FCC, where required reaction rates can be so high as to induce diffusion limitations in a catalyst. Optimized porosity provides for rapid diffusion of feed and product molecules to and from the active sites. DMS products also contain stabilized, mesoporous, alumina-rich matrix components for additional bottoms cracking. Further cracking of ring structures can occur in secondary steps on these acidic surfaces. By microscopy, these materials can appear to be occluded by the primary zeolite over-layer. Yet, laboratory and commercial data show that the results represent breakthroughs in both bottoms cracking, either at constant conversion or constant coke, as well as in coke selectivity [1,2]. The matrix components of these catalysts are therefore also important and interesting materials. Alumina-based active matrix materials are commonly employed in FCC catalyst preparation, as well as in other catalytic applications. Aluminas with relatively large aggregate particle size, such as gibbsite, bayerite, and nondispersible boehmite, as well as dispersible boehmites with smaller aggregate particle size, have all been used as matrices in FCC catalysts. Upon calcination these yield transition aluminas, the more stable of which have either a defect spinel structure or a distorted version of this structure [5]. The 7, 6, and 0 forms derived from boehmite each have cubic close packing of oxygen anions. This cubic packing in the oxygen lattice is preserved during sintering [6], up until its conversion to hexagonal close packing in tx-A1203. This last transition of course leads to the final loss of surface area and the generation of an exotherm in the DTA profile [6]. Soled [7] considered the defect spinel model for ')'-A1203in detail, proposing that surface hydroxyls be accounted for with the overall stoichiometry A12.sO3.5(OH)0.5. Soled also proposed a coalescence mechanism for surface area loss involving dehydroxylation and annihilation of cationic and anionic vacancies [7]. Tucker [8] has found that ctA1203 nucleates at the exteriors and necks between alumina crystallites, confirming the mechanism. Burtin et al [6] elaborated on the dehydroxylation model, and were able to construct a mathematical model that simulated conversion to ~-A1203.

259

Stabilization of 7-A1203 has been pursued for high temperature catalytic applications [9] for some time. Engelhard patents filed in 1973 [10], for example, were part of a developing technology demonstrating surface area and structural stabilization against ot-A1203 transition at 1200 ~ C. Various combinations of group IVB and VIB, and rare earths (especially ceria) were claimed. Improved catalytic performance with rare earth reported by others [11] may have had similar origins. In the open literature, Schaper et al [ 12] were apparently first to report the stabilization of y-A1203 by lanthanum doping. Additional reports confirm stabilization of alumina by Zr, Ca, Th, [13], Ce 3+ [14], Ce 4§ [15], Ba, Sr, and Si [16, 9, 17, 18]. Intimate surface contact between these oxides and alumina in general [ 10, 9], or grafting onto surface hydroxyls in the case of Si in particular [17], are thought to be important in stabilization. These ideas relate back consistently to Soled's surface dehydroxylation model for 7-A1203 [7]. Whether by physical separation, titration or other modification of surface hydroxyls, inhibition of condensation of adjacent particle surface hydroxyls appears to be important in the stabilization process. Most recently [9], synthesis of hexaaluminate phases with many of the same doping species now in the bulk have led to improved surface areas for very high temperature applications. This suggests the best results for Si stabilization of 7-A1203 will be obtained when Si is contained in the bulk of the 7-alumina spinel structure. The purpose of this paper is to report characterization results obtained for model materials related to the DMS matrix component of our FCC catalysts. Since these FCC catalysts contain zeolite Y bonded directly to the DMS, the actual matrix cannot be isolated from the catalyst, and model materials have been studied instead. The precursors of the stabilized alumina matrices of the FCC catalysts are prepared by synthesizing, through hydrothermal treatment and crystallization, a non-zeolitic modified silica-alumina material containing DMS crystallites and/or silica-stabilized y-A1203. Precursor materials may initially have the generic formula (A1203*2SiO2*2H20), and contain an excess of silica with respect to typical A1-Si spinels or other crystalline reaction products such as mullite [19]. In these cases, the reactions can also form phase-segregated SiO2. In general, synthesis of silica-aluminas at stoichiometric or other ratios from alkoxides and salts of Si and A1 can be performed [20, 21], wherein the amount of surplus silica can vary, and these studies have provided great insight into the mechanisms of alumina phase formation. Controversy had arisen in the ceramics literature [19,22,23,24,25,26] as to how much SiO2, if any, can be incorporated into 7-alumina prepared from aluminum disilicate precursors. It was initially believed that a silica-free gamma alumina [23,27,28,29] was formed. Later results obtained after

260

extracting phase-segregated SiO2 revealed that silica was in fact present, although the TEM/EDS data of Sonuparlak [29] is somewhat of an exception. These authors concluded that SiO2 was less than 10 wt% of the ~,-alumina, if any SiO2 were present at all. Chemical compositions of 3:2 A1203:SIO2 [22,30] or gamma alumina containing 8% SiO2 [24] have also been proposed. Coincidental crystallization of an unidentified zeolitic phase has complicated the extraction approach for some [22,31], but not for other researchers [32]. Interpretations of NMR data obtained on y-alumina differed [23,33,34], but it now appears that NMR and TEM/EDS data are in fact consistent with the reported 8% SiO2 in y-A1203 [34,29,26,21 ]. Another valued alumina-silica phase that has been extensively studied in the ceramics literature is mullite, and analogies to these findings may be useful. Mullite can be viewed as solid solution, where the stoichiometry can vary between 3:2 and 2:1 A1203:SIO2 and the magnitude of the first lattice parameter varies smoothly with the alumina content [35]. Immature mullite crystallites made from single-phase sol-gels are initially tetragonal, but transform to the orthorhombic form as the crystallites mature and enrich in Si [36]. Recent papers having thorough discussions on the kinetics of mullite formation [37] and phase characterization by TEM [38] are available to the interested reader. The literature has thus shown that a variety of products may form in the silica-alumina system. Comparison of reaction products made with sol-gel and aluminum disilicate starting materials has shed light on which of these form as a practical matter and why [21, 39]. Silica-aluminas made by slow hydrolysis are intimately mixed, contain many Si-O-A1 linkages, and as a result crystallize alumina-rich mullite directly upon thermal treatments at 980~ [21,39]. On the other hand, poorly mixed A1203 and SiO2 made from granular particles or sols initially yield Si-free 7-A1203 and later Gt- A1203, with mullite forming only at temperatures near 1350 ~ C. Silica-aluminas with intermediate mixing, such as found in layered structures, yield intermediate results. ?-Alumina containing 78 wt% SiO2 is formed simultaneously with Al-rich mullite at 980 ~ C, and aA1203 is never formed [21,39].

2

EXPERIMENTAL

2.1 Material Preparation Matrix precursor materials of two different particle size distributions were used to prepare the model FCC catalyst matrix we have characterized, each precursor having the general formula (A1203*2SiO2*2H20). These precursors do not contain the desired phases and surface area of the active DMS matrix however. Alumina-based catalysts generally require activation before use, and

261

we have followed the procedures of Speronello [40] to thermally and chemically activate, and then characterize our materials. The hydrated aluminum disilicate precursors were first thermally treated in cordierite trays and pre-heated electric furnaces, at a series of temperatures between 980-1300" C, typically for 2-4 hours. In some cases, rotary calcination has also been used. Extractions [40,41] were subsequently done to remove the silica phase and reveal the porosity and surface area of the aluminabased active matrix. Two extractions using 30 grams of solids per 100 grams of 25 wt% NaOH solution were used. Extractions were carried out with stirring at 80 ~ C for one hour, and followed by washing with two volumes of water per volume of NaOH solution. Ammonium exchanges on the solids after extraction employed 1 g NH4NO3:1 g solids: 2 g H20 at 82 ~ C and pH=3, or 66 ~ C and pH=5. A final washing with water and drying gave the activated model alumina matrix samples. Eleven comparative samples of zeolite-free active alumina matrix were also prepared in microsphere form. These samples contained either high or low surface area boehmite, gibbsite, bayerite, flash-calcined gibbsite, or Sistabilized y-A1203. Each of these aluminas were spray dried at 40 wt% loading, with the balance being 20% colloidal SiO2 binder and 40% hydrous clay filler. A control of 20% silica, 80% clay filler was also made. These samples were washed and exchanged as above to remove sodium, steamed at 1350 ~ 1450 ~ or 1600 ~ F, (732 ~ 788 ~ 871 ~ C) and then analyzed for surface area and cracking activity. 2.2

Characterization Methods

Particle sizes for the hydrated matrix precursors were determined by sedimentation, pore volumes for the activated matrix materials were determined by both mercury and nitrogen porosimetry [40], and chemical compositions by XRF. An Autopore IV mercury porosimeter was used with contact angles of 140 ~ An ASAP 2400 was used with the BJH calculation method to determine nitrogen pore size distributions. A Philips PW1877 diffraction system was used with CuK~ radiation, with generator settings of 45kV and 40mA. Crystallite sizes were calculated using the Scherrer equation for the (110) reflection. We conducted secondary electron image analyses on a JEOL JSM-6500F Schottky Field Emission Scanning Electron Microscope (FESEM) at 5-10kV and a 10ram working distance. The powders were mounted on a stub and coated with 5-10 nm of platinum using a Denton DSM-5A sputtering unit in a DV-502A vacuum evaporator for viewing in the FESEM. Transmission electron imaging was conducted on a JEOL 2010 STEM. Samples for TEM analysis were prepared by mounting in Buehler Epofix epoxy/hardener and

262

ultra-microtoming to 90-150 nm thickness using a Leica Ultracut microtome. Some activated ~/-alumina samples were sprinkled directly on a copper grid for TEM/EDS. The EDS analyses were conducted using the Kc~ peaks at 25kV (FE-SEM) and at 200kV (TEM) on PGT SPIRIT workstations equipped with 50 mm 2 light element (C, O, and N) germanium detectors. The Cliff-Lorimer standardless method was used to estimate compositions, using library calibration factors provided by the manufacturer. Diffuse reflectance Fourier-transform infrared (DRIFT) spectra were recorded on a Perkin-Elmer Paragon PC1000 spectrometer equipped with a MCT detector and a Spectra-Tech diffuse reflectance high temperature chamber. Samples dehydrated at 450 ~ C were cooled to room temperature for pyridine adsorption. Physisorbed pyridine was removed by heating to 180 ~ C. A variation of the method of Alerasool et al [42] has been used to determine matrix activity on zeolitic FCC catalysts. The first modification was that USY containing FCC catalysts were impregnated with 50 wt% HNO3 solutions, then dried and calcined, in order to destabilize the zeolite. The samples were then steamed at 1600 ~ F [42], after which no zeolite was detected. The second modification was that microactivity tests (MATs) run in duplicate were used to measure the matrix cracking activity directly. Most of the MATs were run on gas oil feed at 910 ~ F (488 ~ C), 15 WHSV, and 48 sec oil delivery time. Other runs were done on a second gas oil with slightly higher CCR at 970 ~ F (521 ~ C), 24 WHSV and 30 sec oil delivery time, yielding equivalent conversion results but slightly higher coke selectivity. The conversion X was defined on the basis of the wt% yields of 430 ~ F (221~ C) and lower boiling products plus coke, and the effective second order activity for plug flow was defined as (X/(1-X)).

3

RESULTS AND DISCUSSION

3.1 Development and control of macroporosity. Our guiding hypothesis was that further improvements in SCT could be obtained by optimizing the pore volume distribution and matrix dispersion of SCT FCC catalysts [1]. The particle size of the active matrix was reduced (Figure 1) by an order of magnitude relative to conventional alumina matrix materials. At constant wt% loading of matrix, decreasing particle size increases both the number of particles and the number of particles per unit volume. This enhances the distribution of the matrix throughout the FCC microsphere, which reduces the diffusion path length in the microsphere to the nearest matrix particle. Reduced particle size precursors also reduce the diffusion path length inside the matrix, which could hypothetically be beneficial due to the cracking activity and mesoporous nature of an active

263

Figure 1. Comparison of sedimentation particle size distributions for (,,) DMS precursor particles and (A) and a conventional matrix alumina sample as a control.

matrix in general. Considering the platelet morphology of the DMS precursors, the effective diffusion path length within the active matrix domain may have been reduced by as much as two orders of magnitude. If transport processes become important on these length scales in SCT FCC, these changes can increase matrix effectiveness without changing acidity, matrix surface area, coke or gas. Simultaneously, macroporosity control processes were developed to tailor macroporosity in the finished catalyst. Figure 2 shows an SEM illustrating the "house-of-cards" structure of the heat-treated DMS precursor aggregate. The platelet morphology of the original precursor particles is still evident in the structure, and this morphology provides increased macropore surface area and improved accessibility to these surfaces. In the full DMS FCC catalyst, the zeolitic phase is bonded to the surfaces of these macropores. Also shown is an image of the coarse particle size activated alumina sample referred to in Figure 1. While the DMS aggregate is similar in particle size to other common granular FCC matrix aluminas, the DMS aggregate structure contains substantial void volume and macropore surface area not present in alternative materials. The upper part of Figure 3 compares the macroporosity results obtained on three DMS precursors with systematically varying degrees of DMS crystallinity (discussed below) against a sample of the coarse alumina control.

265

Each of the samples in the figure was activated, but did not contain zeolite Y. The generally higher pore volume and wider diameter macropores for the small particle size DMS precursors versus the coarse control are the result of the "house-of-cards" morphology. This matrix structure is robust enough to survive activation and other processing, and so translates to the FCC catalyst. As an example of the robusmess and process technology for macroporosity control, the lower part of Figure 3 shows results obtained with a non-activated sample that originally had a bimodal distribution of macropore widths. Both total pore volume and modal pore diameters were manipulated, suggesting some malleability in the aggregates. Presuming the malleability is a function of the hydrothermal treatment conditions, the macroporosity shift in the activated DMS in the upper part of Figure 3 may well be due to compression of the aggregates during Hg intrusion.

3.2 Development and Control of Mesoporosity. Heat-treated matrix precursors were also examined by XRD before and after activation. Figure 4 gives an example where the difference in thermal treatment was 50 ~ F (28 ~ C). A sample treated at 2000 ~ F (1095 ~ C) contains a non-zeolitic crystalline silica-alumina (DMS crystallites), a spinel phase indistinguishable from 7-A1203, amorphous SiO2 and a trace of anatase. The coexistence of the DMS crystallites and the 7-A1203-1ike phase in Figure 4 is further supported by literature data for silica-aluminas that have an intermediate level of mixing [21 ]. As the crystallization temperature increased, the amount of the crystalline phase, which is indicated by the sharp diffraction lines, increased while the amounts of SiO2, and 7-A1203 (peak positions are given by the legend markers) decreased, since the latter were being consumed by the formation of the crystalline DMS phase. In the most crystalline sample of the figure, nearly all of the y-A1203 has been consumed. The increase in crystallinity is accompanied by an increase in crystallite size, as would be expected. Lee et al [38], among others, noted a rapid growth in crystallite size upon the decomposition of the y-alumina phase at about 2200 ~ F in their materials. The samples of Figure 4 gave a DMS crystallite size by line broadening in the 14 nm to 19 nm range, see Figure 5. Because the inherent defects in the structure of 7-A1203 and its modifications will also contribute to line broadening in XRD, we have not determined the effective crystallite size of the stabilized 7-alumina in this way. Others [24] reported a particle size of 5-8 nm by XRD. TEM results for our materials will be presented further below. Surface areas, pore size distributions, and chemical compositions were determined after activation for a series of model DMS samples of varying crystallinity, and these are reported in Figure 6. The A1203/SIO2 ratio trended

266

Figure 4. XRD from thermally treated but not extracted samples indicate stabilized ~/-alumina is consumed during DMS crystallization, and that the crystalline and 7-A1203 phases coexist.

Figure 5. XRD line broadening by DMS crystallites in the (l 10) reflection.

267

Figure 6. Correlation of activated model DMS (O) t-plot surface area, (A) 2V/A pore diameter, and (m) A1203/SIO2 molar ratio with DMS crystallinity by XRD.

Figure 7. BJH nitrogen pore size distributions on activated model DMS at (+) 5%, (A) 27%, (x) 45%, (D) 50%, (O, and 0) 56% crystallinity.

268

to 1.5 for the high crystallinity material. Less crystalline material contained more alumina, in directional agreement with literature results [21 ]. This shifts unit cell size [35] and may possibly influence acidity in a manner analogous to what happens with zeolite Y. The silica-stabilized gamma alumina samples gave an A1203/SIO2 ratio of about 2.1 however, which is similar to some reports in the literature [22, 44], but in poor agreement with others [19,26,27]. Contamination by sodalite has been found in the most recent stabilized 7-A1203 (5% crystallinity) samples however, so this composition may overestimate the SiO2 content. This problem is discussed further in section 3.3. Activated surface areas between about 30 and 190 m 2 g-~ and volumeaveraged pore diameters between 50 and 200 A are also reported in Figure 6. Although the relationships are not quantitative, these surface areas and pore diameters are reasonably related to particle sizes by microscopy, suggesting that the fundamental particles are nonporous, and the pores represent voids between crystallites. The changes in surface area and pore diameter at intermediate crystallinity are due partly to crystallite size and partly to the relative DMS crystallite and stabilized 7-A1203 contents of the matrix. For example, Figure 4 showed intermediate materials contained both the semicrystalline and the fully crystalline components. Accordingly, the N2 pore size distributions of Figure 7 are somewhat broader at intermediate crystallinity. The legend in the upper part of Figure 3 had indicated the DMS crystallite integrated diffraction intensities relative to a well-crystallized, pure reference standard. The cause of the shift in mesopore diameter is now appreciated as being due to growth in the amount and size of the DMS crystallites. Mercury and nitrogen mesoporosity distributions generally agree in this area. Electron microscopy was used to characterize the activated materials as well. We found well-crystallized DMS in general to have a cylindrical morphology (Figure 8) with diameters in the 40-90 nm range. We note however, that the Pt coating may be thick enough in these images to affect the apparent particle size. Employing these apparent diameters and aspect ratios in some cases gave calculated surface areas in reasonable agreement with observed values, implying that the surface area is associated only with the exterior of the crystallites. More often however, surface area estimates were significantly below the observed values, and this can be attributed to the influence of the Pt coating on the apparent crystallite size. The external surface area interpretation is further supported by pore size distribution data, which shows little porosity below 100 A for well-crystallized samples (Figure 7). Samples exhibiting lower intensity, broader diffraction lines were found to have somewhat smaller crystallite diameter, but even more significantly reduced L/D aspect ratio (Figure 8, bottom). The disparity between the calculated and observed surface area was large in this latter case. Some of the discrepancy may again be due to the Pt coating used to reduce charging in

269

Figure 8. Above, FE-SEM of well-crystallized, activated model DMS with cylindrical crystal habit. Below, activated model DMS of 27% crystalli',ity, having smaller crystallites. Arrows to the right of each image are 1 gm.

270

SEM. Imperfect agreement is also in accord with diffraction (Figure 4) and pore size distribution results (Figure 7), suggesting that significant levels of the stabilized y-A1203 coexist with small crystallite size DMS. For example, the material shown in the lower part of Figure 8 with 2.75 times higher magnification has a BET surface area of 109 m 2 g-l, which implies a crystallite size of about 18 nm at an aspect ratio of 1. The observed value is roughly 50 nm, nearly half of which could potentially be due to Pt. The pore size distribution obtained for this sample (27% crystallinity in Figure 7) contains significant porosity in the narrower ranges and is almost bi-modal in nature, as expected when diffraction shows that stabilized y-A1203 and small crystallite DMS coexist. It has not been possible to deconvolute these factors quantitatively, but the results are directionally self-consistent once the occurrence of mixtures is considered. TEM has also been used to characterize the materials under study, and the images in Figure 9 further serve to accentuate the differences in particle size and morphology possible in DMS. The top left comer shows the material with the second to the widest pore size distribution in Figure 7. The DMS crystallites appear less ordered than in Figure 8, and this disorder phenomenon appears to relate to disorder in the original hydrated precursors. A nominal diameter of 40 nm, comparable to but probably smaller than that of Figure 9, is required to give a calculated surface area matching the experimental BET of 35 m 2 g-1. The upper right image begins to reveal the lattice of well-crystallized DMS, confirming its well-ordered, long-range crystalline structure. The lower images in Figure 9 show moderate and high magnification views of the unmounted, activated, Si-stabilized y-A1203 obtained without microtoming. The particles at the platelet edge appear to be about 5 nm in diameter and have an aspect ratio near 1. The surface area calculated for this particle size, using a true density of 3.2 g cm -3, is 375 m 2 g-l, which is much higher than what we have observed (Figure 6), non-representative of typical 7A1203 [7], but comparable to what has been reported by others [44]. Okada et al [24] on the other hand reported a TEM particle size "as large as 10 nm." A 10 nm particle size corresponds to 188 m 2 g-l, which is essentially the same as the surface area we have observed, and provides a more sound rationalization of the observed 5 nm modal pore diameter for the present materials (Figure 7). Since the samples were not microtomed, it is unlikely that sectioning has reduced the apparent particle diameter. It is possible that there is sampling bias however, since observations were made preferentially at the edge of an activated platelet. The most significant issue affecting the 7-alumina characterization is contamination by hydroxysodalite however, which is confirmed to have crystallized during the most recent activations of the 7alumina precursors. This will be discussed in more detail in the next section.

271

Figure 9. Microtomed section TEM after activation of (above) well crystallized model DMS treated at 2288 ~ F (left, under focused) or treated at 2350 ~ F (fight). Below, TEM of unmounted, activated, stabilized ~/-A1203. DMS crystallite and stabilized )'-A1203 lattice fringes are visible in the two higher magnification micrographs on the fight.

272

For the moment we note that the surface area and particle size data discrepancies require some 50% sodalite be present for self-consistency, and that the contaminant is readily recognized in microscopy and not present in Figure 9.

3.3 Si content of y-Al~O3-like phase. The simple fact that (~-A1203 is not observed when aluminum disilicate precursors with intimate or intermediate mixing are thermally treated is nearly convincing, in and of itself, that the ")'-A1203 -like phase formed as a metastable intermediate during DMS crystallization in fact contains significant silica. On the other hand, this lone argument does not completely exclude the possibility that the T-A1203 phase is silica-free and that the Si diffuses into it at higher temperatures, forming DMS crystallites instead of c~-A1203. As noted above, there has been a controversy in the literature [ 19-34], but Okada and Otsuka's conclusion [20,21] of about 8 wt% SiO2 in y-alumina appears most reliable. Our purpose here will be to show that our materials are consistent with the label Si-stabilized T-A1203. One hypothesis which explains finding SiO2 in our activated T-A1203 is that the extractions were incompletely removing surplus SiO2. Table 1 reports results for an ultrafine DMS precursor of particle size like in Figure 1, having less than about 5% of the crystalline phase of Figure 4; i.e., nearly pure 7-

Table 1. Repeated extractions of Si-stabilized ~,-A1203precursor of model DMS. Number of Extractions

1

2 3 Liquids analyses

4

A1203,wt% as is

0.06

SiO2, wt% as is

1.8

BET, m2/g

125

86

94

91

XRF Na20, wt%

9

15

15.6

15.4

XRF A1203,wt%

64

60

52

52

XRF SiO2, wt%

22

20

26

25

XRD Sodalite, wt%

8

26

33

38

*SiO2in activated T-A1203,wt%

21

14

20

17

*Na20 in activated y-A1203, wt%

8

14

14

13

0.28

0.25

0.007 0.02 Solids analyses

0.23 0

*Estimated by subtracting contribution from sodalite, if extracted but not exchanged.

273

alumina in a mixture with phase-segregated SiO2. Four extractions were performed sequentially. Because of the very fine particle size, filtration and washing was difficult and complete extract removal could not be assured. More than 90% of the liquids were recovered as filtrate however. Regardless, the liquid analyses showed that most of the SiO2 was dissolved in the first extraction step, while significant SiO2 remained in the solids. Subsequent extractions actually began to dissolve A1 much more selectively than Si. This agreed with earlier data associated with Figure 6, which had confirmed that two extractions were required to fully develop surface area and remove phasesegregated silica. XRD scans obtained later on these solids unfortunately showed the presence of substantial amounts of the zeolite hydroxysodalite. The surface areas were also rather low and inconsistent with circa 190 m 2 g-1 obtained earlier (Figure 6). S i O 2 increasing on the solids is consistent with crystallization of sodalite, and perhaps implies dissolution of SiO2 from the glass beakers during the extraction process. Extractions repeated on other precursors also gave high levels of sodalite. Although this was not apparently such a significant problem earlier, the samples do not exist to test for sodalite contamination and so there is no proof that the 7-alumina samples of Figure 6 were not at least somewhat contaminated with sodalite. Others have obtained higher surface area than 190 m 2 g-~, and without reporting the reason, shown that extended extraction times reduce surface area and increase SiO2 content of the solids [44]. Some investigators have explicitly reported a zeolitic phase is formed during extraction [22,31] while others have reported none is formed [32]. We have not been able to control the same inconsistency within our own results at the present time. In order to derive useful information from these results, five sodalite diffractions above 24 ~ and a sodalite standard estimated to be of about 90% purity (Figure 10) were used to estimate the sodalite contents in the samples of Table 1. Neglecting differences in volatile matter between the 7-alumina and sodalite phases, any non-Si-A1-Na constituents, and assuming an Na20: A1203: 2 SiO2 stoichiometry for sodalite, we were able to estimate and subtract the SiO2 and Na20 contributions due to the sodalite in the mixture. The results in Table 1 estimate that the extracted but not exchanged Sistabilized "y-A1203 contains roughly 14-21 wt% SiO2 and 8-14 wt% Na20. By comparison, the extracted precursors of Figure 6 with 5% DMS crystallinity gave 8 and 9 wt% Na20 after washing with water and about 19.6 wt% SiO2. A modest amount of the SiO2 and Na20 may be due to entrained extract in the present samples, but it seems unlikely that all of the SiO2 is an artifact of unaccounted for zeolite and/or poor washing. For example, the extracted products would each have to be roughly 50% sodalite to account for the case. One could then further rationalize the discrepancies between the 375 m 2 g-I

274

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estimated for 5 nm stabilized alumina crystallites (Figure 9), and the low surface areas of Figure 6 as being due to universal contamination by 50 wt% sodalite. Thus, the results are still not satisfying. Solid-state NMR and IR structure band data were independently obtained, but these results were not conclusive. Other attempts made to determine the SiO2 content of the Si-stabilized y-A1203 in DMS were by EDS, both in SEM and TEM. The contaminating zeolites have a tendency to crystallize in micronsized crystals with distinct morphology. This suggested that we could perhaps employ micron-sized electron beam analysis volumes carefully, avoiding the sodalite and analyzing just the ~-alumina phase. Others have used the same method (absent a contaminating zeolite) to show, seemingly convincingly, that there could not be [29], or was certainly some [24], or a great deal [30] of SiO2 present. Lee et al [38] declined to employ the technique, due to the tradeoffs between spot size and secondary phase contributions. Our own SEM/EDS data on micron sized, Pt-coated samples were also not convincing, as we could not assure that cross-contamination was absent. Instead, electron beam analysis volume was minimized using TEM sections 100 nm thick and a spot size of 15 nm. Figure 11 is a low magnification view showing locations for spot analysis. A portion of the raw EDS spectra are also presented in Figure 11. The model DMS sample was derived from an ultrafine precursor (Figure 1) containing only about 5 wt% DMS crystals by XRD before activation, but sodalite content was not measured.

276

Table 2. Chemical compositions estimated by EDS for the activated model matrix ,/-alumina phase on the spots of Figure 11. Results reported as wt% oxide or A1203/SIO2 molar ratio. TEM Spot 1 2 3 4 5 6 7 8

SiO2 17.6 17.8 19.2 28.2 67.4 4.5 39.8 0.6

A1203 76.6 75.7 75.1 62.9 30.7 1.9 50.3 23.7

TiO2 1.1 1.5 1.0 1.2 0.1 93.5 0.3 71.1

Na20 0 0.7 0.2 2.7 0.1 0 4.6 0

Fe203 4.1 3.6 4.3 2.8 0.1 0.1 1.2 4.6

CaO 0.6 0.7 0.1 2.4 1.6 0.1 3.7 0

A1203/SIO2 2.6 2.5 2.3 1.3 0.3 0.2 0.7 23

The numerical results associated with Figure 11 are summarized in Table 2, and these show significantly different elemental compositions in the chosen areas. By comparison of the EDS and XRF compositional results, or the contrast in the electron transmission results of Figure 11, spots 1, 2, and 3 appear to be most representative of the high surface area stabilized DMS alumina. The agreement between these spots and the XRF results in Figure 6 or Table 1 is generally good, except that the iron levels found by EDS are significantly higher than the quantitative bulk analysis by XRF. Since the EDS results are only semi-quantitative, the actual iron levels are probably lower than indicated in Table 2. These data are also more consistent with the Sicontaining y-alumina described by Okada and co-workers [20,21,24,26] than with Si-free gamma alumina, although the wt% SiO2 results are higher in our case than reported by Okada. The other spots in Figure 11 were chosen for analysis because of their electron transmission contrast versus the bulk of the material. Spectra for these other spots showed much different A1203/SIO2 ratio than the bulk material. The cation content and contrast of spots 4 and 7 suggest a non-exchangeable dense phase, but the A1203/SIO2 ratios do not match well the 0.5 value expected for sodalite. Spot 5 was very dark in Figure 11 but higher in Si content than sodalite. The ion exchange procedures significantly degrade x-ray diffraction by sodalite however, so exchangeable cation content can not be used to differentiate between phases. While these phases have not been identified, they would appear to be non-representative of the bulk. Finally, TiO2 represents only a tiny fraction of the DMS overall according to XRF, but is the dominant component of spots 6 and 8. These spots are therefore also isolated impurities whose composition is not representative of bulk values.

277

Since the activated surface area of this sample was relatively high, the analyses for representative spots agreed well with bulk values, and repeated extraction data indicated two extractions had been sufficient, the most consistent interpretation is Si-stabilized 7-A1203, speculatively containing about 8-20 wt% of SiO2. If Si-free ]t-A1203 was prevalent instead and the found SiO2 were an artifact of micron scale sodalite contamination, SiO2 should have been localized to certain spots like the Na and Ca, instead of being well dispersed. The contrary hypothesis, while consistent with much of our data, is not supported by the TEM/EDS results.

3.4 Acidity, stability and cracking activity The acidity of the activated model DMS samples was measured using FTIR with pyridine. We found that the materials have some Br6nsted acidity, but the acidity was mainly of the Lewis type. Since Br6nsted acidity is associated with Si-OH-A1 groups and ~-A1203 does not have Br6nsted acidity, finding these sites on the y-A1203-1ike DMS phase is additional evidence for the presence of surface silicon and structural stabilization. The IR spectrum of pyridine adsorbed on the 5% crystallinity sample of Figures 6, 7 and 11 is therefore presented in Figure 12. The Lewis acidity was subdivided into three strengths: strong, medium and weak. These were differentiated by temperature of desorption and quantified by curve fitting of the bands at ~1450 cm -~. As shown in Table 3,

0.2 L

0.15

=~

o.1

o

0.05

0

-0.05 1700

I

I

a

1600

I

1500

i

1400

Wavenumbers (cm-1)

Figure 12. Br6nsted acid sites found at 1546 cm -1 by pyridine IR on the 5% crystallinity DMS also confirm the existence of Si on DMS ~'-A1203.

278

Table 3. Soda, surface area, A1203/SIO2 molar ratio and acidity data for activated model DMS matrix materials, versus percent crystallinity by integrated powder diffraction areas. Crystallinity, % Na20, wt% BET, m 2 g-1 A1203/SIO2, m/m

56 0.055 29 1.45

Br6nsted Strong Lewis Medium Lewis Weak Lewis Total Lewis

0 34 123 5 162

Total acidity/BET

5.6

56 0.09 35 1.49

50 45 0.13 0.09 59 71 1.51 1.57 Acidity, gmol g -1 3 8 10 60 91 95 99 190 300 38 14 20 197 295 415 Total Acid site density, pmol 5.7 5.1 6.0

31 0.09 104 1.64

5 0.17 199 1.99

32 82 335 27 444 m -2 4.6

40

362 2.0

most of the crystalline DMS sites were moderate in strength; some were strong and only a few weak sites were found. The total Lewis acidities ranged from 162 to 444 gmol/g and, with the significant exception of the 5% crystallinity sample, the total acid site density correlated well with surface area, giving nearly constant acid site density. This suggests that the expected systematic changes in Si/A1 ratio and unit cell size of the crystalline DMS crystallites (Figure 6) and analogous solid solutions [35] are not influencing surface acidity. With regard to the Si-stabilized 7-A1203, although the acidity data were reproduced, MAT results reported below imply that the acid site densities of these materials are actually quite similar to the crystalline DMS and other matrix aluminas. A zeolitic FCC catalyst containing DMS was found to have 308 gmol/g of Lewis acidity. Acid site distributions are important in FCC catalyst performance, but when we compared the performance of DMS-containing catalysts to other FCC catalysts with equal Lewis/Br6nsted or ZSA/MSA ratios at constant steamed unit cell size, significantly improved performance is obtained for the former [ 1,2]. The improved performance is believed to derive not from acid site distribution per se, but instead from the unique macroporous zeolite-on-matrix morphology, which alters the sequence of sites contacted in cracking, reduces active site occlusion, minimizes diffusion path lengths and maximizes diffusivity. The macroporosity and diffusion path length features are related to the choice of DMS precursor particle size and creation of aggregates of these particles (Figures 1-3). Zeolite is bonded to the external surfaces of the DMS platelets in the finished FCC catalysts. The platelets

279 themselves are composed of DMS crystallites of controlled particle size and/or silica-stabilized 7-A1203. Steam stability was assessed for the activated model DMS materials, and compared to commonly available matrix aluminas, the latter being spray dried at 40 wt% loading with binder and clay filler. The results showed that the surface area retention for the various sol-bound aluminas averaged 50% at 1450" F and 37% at 1600 ~ F. By comparison, three model DMS activated matrices gave 74-90% retention at 1500 ~ F and one test gave 82% surface area maintenance versus fresh after steaming at 1600 ~ F. While much of the matrix surface area loss in the comparative cases was in fact due to silica sol binder sintering, the stabilities of the steamed DMS materials are independently excellent. As noted above, silica stabilization of y-AI203 is the most effective option [16, 17, 9], presumably because of elimination of surface aluminum hydroxyls [7,6,9,17] and/or because of the higher charge on the Si cation [6]. The surface silanols anticipated via Beguin's work [17] were in fact found in our DRIFT spectra at 3736 and 3724 cm ~, on both crystalline and Si-stabilized y-A1203 samples of activated model DMS. The fact that Si-O-A1 linkages are very strong is probably also important. Similarly, phase diagrams for this composition range imply the crystallites are highly stable matrix materials, yet they are surprisingly flexible in terms of surface area, pore diameter and acidity (Figures 6,7; Table 3). Cracking activity has also been determined for the matrix technologies described above. As expected from their compositions, acidities and surface areas, the Si-stabilized y-A1203 and crystallites of DMS are found to be active cracking catalysts. What may not be so readily appreciated is how similar matrix materials can be in terms of specific activity and coke selectivity. In what follows, microactivity test (MAT) results for two sets of catalysts are presented. The first is the same set of eleven zeolite-free, sol-bound active alumina matrix materials described above. The other set of samples was prepared from a range of commercially-available FCC catalysts comprising several manufacturing technologies, with the goal of characterizing their matrix cracking activities. Following Alerasool et al [42] in part, USY-containing FCC catalysts were impregnated with HNO3 and steamed at 1600 ~ F. MATs run in duplicate were then used to measure the cracking activity of the two sets o f samples. The results plotted against BET surface area in the upper part of Figure 13 show that the cracking activities of common FCC alumina matrix materials are very similar, after accounting for differences in surface area. The only exceptions we have found to the trend are the FCC sample in the figure steamed at 1600 ~ F with 126 m 2 g-l, which had rather high sodium content, and the zeolite-free gibbsite-derived alumina steamed at 1350 ~ F and giving about

280

Figure 13. Second order MAT activity (above) and coke selectivity (below) of zeolitefree activated alumina matrices after (+) 1350 ~ F, (x) 1450 ~ F or (-~) 1600 ~ F steaming, and of (0) FCC catalysts after acidification and 1600 ~ F steaming. MATs run at 910 ~ F and 48 sec oil delivery time, with X representing 430 ~ F-based wt% conversion. Other MATs (A) were run on acidified FCC catalysts after 1600 ~ F steaming at 970 ~ F and 30 see oil delivery time with a similar gas oil, giving equivalent activities but higher coke selectivity than before. Overall, specific cracking activities and coke selectivities are remarkably similar.

281

1.35 activity. In this latter case, micropores can be formed [5] and, absent another explanation, we can only speculate that capillary condensation of gas oil may have increased the specific MAT activity. The activity of DMS-related materials followed the general trend in the figure, and if they materially differed, their specific activities were biased slightly lower. For the most part however, bottoms upgrading aluminas appear to have equivalent specific activity. Plotted in the lower part of Figure 13 are coke selectivities from most of the MAT runs. Two trends are found, with the 970 ~ F MATs producing systematically higher coke selectivity, which can be attributed to differences in the test methods. Otherwise, the data are remarkably consistent, surprisingly showing no high coke selectivity aluminas were found, but perhaps identifying a few activated aluminas lower in coke selectivity. Among these are the gibbsite-derived alumina samples and the sol-bound clay control sample after mild steaming. These seemingly lower coke selectivity aluminas have not been unusually successful in FCC overall however, and can be dismissed. Testing matrix without zeolite present perhaps oversimplifies the real situation, eliminating any possible zeolite-matrix synergies. It is therefore probably na'fve to expect matrix-only coke selectivity differences to project linearly through to the zeolitic FCC catalysts. It is perhaps for these and other reasons that the DMS-related samples in these tests did not show materially different coke selectivity. Thus, the alternative hypothesis that the improvement in DMS bottoms cracking activity or coke yield at constant conversion is caused either by a systematically higher non-zeolitic matrix acid site density, strength, specific cracking activity, or by dramatically lower non-zeolitic matrix coke selectivity, is not supported by the available data. It remains possible that the detailed selectivities of alternative active matrix materials can differ, but the correlation of Alerasool [42] and our own experience (prior to the advent of DMS technology) suggested that their activity for bottoms upgrading at constant matrix surface area would be comparable. We were therefore surprised to find [1,2] that the materials now termed DMS in fact delivered a discontinuity in bottoms upgrading at constant conversion, zeolite unit cell size and ZSA/MSA. The data of Figure 13 and Alerasool [42] show that MSA is a de facto measure of matrix activity, and it is well known that unit cell size and ZSA determine zeolite activity and gasoline selectivity. Comparison of cracking selectivities while holding ZSA/MSA and steamed UCS constant therefore excludes their systematic influence on performance, leaving only true performance differences. We concluded [1,2] therefore that it was not the matrix chemistry per se which yielded the breakthrough, but rather the unique zeolite-on-matrix morphology of the full DMS FCC catalysts that gave rise to performance improvement.

282

Figure 14 presents a final SEM of the DMS FCC catalyst containing this unique structure, where the matrix precursor was in the form of very thin, broad platelets. Planes of polycrystalline zeolite Y are seen on the two external surfaces of the plates, in morphology reminiscent of molecular sieve membranes. The relatively smaller grains at the center of the planes are wellcrystallized DMS, the large crystallite size of which then generates a low matrix surface area as shown in Figure 6. Catalysts based on such formulations yield ultimately low ZSA/MSA and catalytic coke, together with bottoms upgrading paradoxically comparable to conventional catalysts at the other end of the ZSA/MSA spectrum [ 1,2]. 4

CONCLUSIONS

Model DMS matrix materials were characterized for morphology, structure, composition and acidity after a suitable activation process. DMS crystallites and Si-stabilized T-A1203 could be prepared either with relative purity, or in mixtures. Variation of their relative amounts in mixtures and/or the crystallite size of the DMS led to smooth variations in surface area and mesopore diameter, enabling preparation of FCC catalysts over a wide range of Z/M. The DMS crystallites have a cylindrical crystal habit, the diameter and aspect ratio of which increases with crystallization severity. The DMS crystallites themselves are nonporous, so that their surface area and mesoporosity are associated with the external surfaces and the spaces between the crystallites. The A1203/SIO2 ratio for the well-crystallized DMS is close to 1.5, but approaches 2 to 2.5 at lower relative crystallinity. Despite this significant change in silica-alumina ratio, no convincing change in acid site density and cracking activity was found. TEM/EDS has confirmed that a Si stabilizer is present in the model DMS T-A1203 component, and the results agreed well with the 15-20 wt% SiO2 found by bulk analysis on the same samples. Fundamental particle diameters of 5 nm were determined by TEM, but surface area and porosity data suggest an 8-10 nm particle size may be more representative. The activated surfaces of either phase contain mainly Lewis acidity, and have specific cracking activities and coke selectivities comparable to other common FCC aluminas. Hydrothermal stabilities are excellent. DMS macroporosity is created via a "house-of-cards" aggregate morphology, and this is optimized and controlled by selection of matrix precursor particle size and other process parameters. Ultra-fine precursor platelets reduce the internal and external diffusion path lengths of the active matrix. This helps improve matrix effectiveness without changing MSA, acidity, coke or gas.

283

Figure 14. SEM of a zeolitic DMS FCC catalyst made from well-crystallized DMS precursor platelets with a very high aspect ratio, fractured to reveal the interior of the microsphere. Above, 0.2-0.5 ~tm faujasite crystallites are bonded to the surfaces of the macropore walls. Below, broken platelets reveal that planes of polycrystalline DMS form the interior of a composite material having an overall macroporous zeolite-on-matrix morphology.

284

These and other supporting data suggest that the improved selectivities of the DMS technology in FCC is not due to a change in surface acidity per se, but rather to the novel macroporous zeolite-on-matrix morphology of these FCC catalysts. This unique structure changes the sequence of acid site contacting in cracking in a favorable way, reducing coke selectivity and improving the yield of the primary products of cracking.

References [11

[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]

D. M. Stockwell, presented at 3ra Intl. Conf. on Ref. Proc., AIChE National Meeting, March, 2000. J. B. McLean, and D. M. Stockwell, Nat. Petr. Ref. Assoc. 2001, AM-01-58. J. B. McLean, D. M. Stockwell, W. S. Wieland, and W. S. Winkler, presented at the European Refining Technology Conference, November, 2001. J. B. McLean, W. A. Weber, and D. H. Harris, Nat. Petr. Ref. Assoc. 2003, AM-03-38. K. Wefers and C. Misra, Alcoa Technical Paper 19, (1987). P. Burtin, J. P. Brunelle, M. Pijolat and M. Soustelle, Appl. Catal., 34 (1987) 239. S. Soled, J. Catal., 81 (1983) 257. D. S. Tucker, J. Amer. Ceram. Soc. 68 (1985) C-163. H. Arai and M. Machida, Appl. Catal. A: General, 138 (1996) 161. S.G. Hindin, and J. C. Dettling, US Patents 3,945,946; 3,956,188; 3,993,572 (1976); 4,008,037 (1977). F.J. Sergeys, J. M. Maselli, and M. V. Ernest, US Patent 3,903,020 (1975). H. Schaper, E. B. M. Doesburg and L. L. van Reijen, Appl. Catal., 7 (1983) 211. P. Burtin, J. P. Brunelle, M. Pijolat and M. Soustelle, Appl. Catal., 34 (1987) 225. A. Piras, A. Trovarelli and G. Dolcetti, Appl. Catal. B: Environmental 28 (2000) L77. M. Ozawa and M. Kimura, J. Mat. Sci. Lett. 9 (1990) 291. C.Z. Wan, and J. C. Dettling, U. S. Patent 4,624,940, (1986). B. Beguin, E. Barbowski, and M. Primet, J. Catal. 127 (1991) 595. R.J. Lussier, S. Plecha, C. C. Wear, and G. D. Weatherbee, U. S. Patent 6,451,200 B 1 (2002). G.W. Brindley and M. Nakihara, J. Amer. Ceram. Soc., 42 (1959) 319. K. Okada and N. Otsuka, J. Amer. Ceram. Soc., 69 (1986) 652. K. Okada and N. Otsuka, in Ceramic Transactions, Vol. 6, Mullite and Mullite Matrix Composites, S. Somiya, R. F. Davis and J. A. Pask,, eds. Amer. Ceram. Soc. (1990) 375. K. Chakraborty and D. K. Ghosh, J. Amer. Ceram. Soc., 61 (3-4), (1978) 170. M.W. Brown, K. J. D. MacKenzie, M. E. Bowden, and R. H. Meinhold, J. Amer. Ceram. Soc., 68 (6), (1985) 298. K. Okada, N. Otsuka,, and J. Ossaka, J. Amer. Ceram. Soc., 69 (10) (1986) C-251. J.D. MacKenzie and I. M. W. Brown, J. Amer. Ceram. Soc., 70 (9) (1987) C-222. K. Okada, J. Amer. Ceram. Soc., 70 (9) (1987) C-223. H.J. Percival, J. F. Duncan and P. K. Foster, J. Amer. Ceram. Soc., 57 (1974) 57. S. Mazumdar and B. J. Mukherjee, J. Amer. Ceram. Soc., 66 (1983) 610. B. Sonuparlak, M. Sarikaya, and I. A. Aksay, J. Amer. Ceram. Soc., 70 (1987) 837.

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[30] K. Srikrishna, G. Thomas, R. Martinez, M. P. Corral,, S. de Aza, and J. S. Moya, J. Mater. Sci., 25 (1990) 607. [31 ] K. Chakraborty, and D. K. Ghosh, J. Amer. Ceram. Soc., 72 (1989) 1569. [32] A. Aksay, M. Sarikaya, and B. Sonuparlak, J. Amer. Ceram. Soc., 72 (1989) 1571. [33] J. Sanz, A. Madani, J. M. Serratosa, J. S. Moya, and S. Aza, J. Am. Ceram. Sot., 71 (1988) C-418. [34] J.D. MacKenzie, and I. M. W. Brown, J. Amer. Ceram. Soc., 70 (1987) C-222 [35] W.E. Cameron, Ceramic Bulletin 56 (1977) 1003. [36] D.X. Li, and W. J. Thomson, J. Mater. Res. 6 (1991) 819. [37] A. Gualtieri, M. Bellotto, G. Artioli, and S. M. Clark, Phys Chem Minerals 22 (1995) 215. [38] S. Lee, Y. J. Kim and H. S. Moon, J. Amer. Ceram. Soc. 82 (1999) 2841. [39] J.A. Pask, and A. P. Tomsia, J. Amer. Ceram. Soc. 74 (1991) 2367. [40] B.K. Speronello, U. S. Patent 4,628,042, (1986). [41 ] H. Schoenfelder, and H. Ginsberg, U. S. Patent 2,939,764 (1960). [42] S. Alerasool, P. K. Doolin, and J. F. Hoffman, I&EC Research 34 (1995) 434. [43] E.L. Moorehead, J. B. McLean, and W. A. Cronkright, in Studies in Surface Science and Catalysis, Vol. 76, Fluid Catalytic Cracking: Science and Technology, J. S. Magee and M. M. Mitchell, Eds., Elsevier (1993) 223. [44] R.J. Lussier and G. J. Surland, U. S. Patent 4,749,672, (1988).

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 9 Elsevier B.V. All rights reserved.

287

Feedstock Effect on FCC Catalyst Stripping Claudia M.L. Alvarenga Baptista and Henrique S. Cerqueira I PETROBRAS R&D Center (CENPES)/Downstream R&D/TFCC Cidade Universitfiria, Quadra 7, 21949-900 Rio de Janeiro, Brazil.

The effect of feedstock on FCC catalyst stripping was investigated based on prototype FCC unit data from three different feedstocks: blend of heavy gas oil from vacuum distillation and delayed coking units, and two different atmospheric residues (paraffinic and with high carbon residue). The higher the feedstock crackability, the higher is the amount of strippable hydrocarbons. The average hydrogen-to-carbon ratio depends on the feedstock and is lower for the paraffinic ATR. The coke on spent catalyst samples was characterized for all feedstocks, indicating a much lower amount of CH2C12 insoluble coke for the heavy atmospheric residue feedstock. The insoluble coke of the heavy ATR was also more alkylated and/or naphthenic than in the other feedstocks. 1. INTRODUCTION The FCC feedstock composition has a direct impact in the product quality and yield profile, affecting the operation of commercial units. In the particular case of resid feedstock processing, many modifications of typical FCC unit hardware are needed [ 1]. The catalyst stripper is a piece of equipment that has gained importance with the processing of more heavier feedstocks in FCC units worldwide and there

I To whom correspondence should be addressed. Fax: +55-21-3865-6626. E-mail: [email protected]

288

are different internals devices to improve its performance [2-4]. Actually, an enhancement of the stripper efficiency will reduce hydrocarbon entrainment to the regenerator where it is burned as coke. For a given feedstock, the higher the stripper efficiency, the lower the amount of hydrocarbons dragged to the regenerator, reducing the coke on catalyst and lowering the dense phase temperature. In the industrial FCC unit, this results in either a higher CTO (higher conversion) or lower demands on the catalyst cooler allowing refiners to process heavier feedstocks. The present paper discusses a large set of data from tests in PETROBRAS' prototype FCC unit with a wide range of feedstocks - a blend of heavy gas oil from vaccum distillation and delayed coking units (HGO + CGO) and two atmospheric residues (ATRs), one much more paraffinic than the other. For each feedstock the unit would naturally stabilize at a different CTO because of thermal balance constraints. In order to compare the feedstocks at iso-CTO, higher catalyst circulations were forced on the heavier feedstocks with the help of a catalyst cooling coil placed in the regenerator. 2. EXPERIMENTAL The experiments were carried out in PETROBRAS' prototype FCC unit, located in S~o Mateus do Sul, State of Paranh, Brazil. This is a heat balanced sideby-side short contact time unit designed by PETROBRAS to process feedstocks with high carbon residue (up to 9wt.%). The feedstock flow rate depends on the feedstock quality and is in the range of 100-200 kg.h -~. The catalyst circulation rate depends on the feedstock and on the amount of heat removed from the regenerator and lies in a range of 500-2000 kg.h -~. The catalyst inventory is 250 kg. The amount of heat removed in the regenerator was modified by changing the air flow rate from 0 to 250 kg.h -1 to the cooler coil. The temperature of the air used to bum the coke in the regenerator was ambient (~35~ for the ATR, 550~ for the gas oil blended with coke gas oil and 670~ for the paraffinic ATR. In the experimental design three distinct catalyst-to-oil ratios (CTO) were considered: the one obtained without removing heat from the regenerator (dense phase temperature close to 710~ and two other levels with lower dense ~hase temperatures. The stripper steam flow rate was constant and equal to 7.5 kg.h-. The hold-up level was 75% of the stripper height, which covered exactly the top stripper baffle and the electrical heating resistances that compensated thermal loss

289

in the stripper region were on. Under these conditions the stripper of the prototype FCC unit will outperform a typical industrial unit one. Stripping steam was injected in two distribution tings in the stripper vessel, half of the total throughput in each one. At least two runs were done for each experimental condition. Reaction temperature was 540~ in all runs. The hydrogen to carbon ratio of the coke was calculated from the flue gas chromatographic analysis and the CTO from the delta coke obtained from carbon content analysis of the spent and regenerated catalysts. Selected operational conditions are summarized in Table 1. Table 1 Operational Conditions. Variable

Run length (h) Reaction temperature (~

value

1 540

Dispersion steam temperature (~

250

Feedstock temperature (~

300"

Riser pressure (kgf. cm 2)

1.6

Lift steam feed (kg.h-l)

4.0

The feedstock temperature for the blend HCO + CGO was 226~

The feedstocks tested had very different properties (Table 2). The parrafinic ATR had the lowest density. The heavy crude ATR had the highest density, asphaltenes content, carbon residue, aromatic content and basic nitrogen, which would indicate a higher tendency to form coke. In order to achieve comparable conversion levels for the different feedstocks, a more active resid catalyst was used in the heavy ATR case [5]. The two catalysts are from the same type and some characteristics are presented in Table 3. The product yield profiles were obtained from a reaction mix sampling (RMS) port located in the converter exit line. The liquid fraction was analyzed by Simulated Distillation with the following cut points: naphtha = 35-221~ LCO = 221-343~ decanted oil - 343+~ Gases were analyzed by gas chromatography,

290

and the C5+ in the gas stream was added to the naphtha yield. The coke yield was calculated from the flue gas chromatographic analysis.

Table 2 Feedstocks Characterization. Analysis

Density @ 20/4~ (g.cm-3)

HGO + CGO

ATR

Paraffinic ATIR

0.9390

0.9595

0.8890

0.6

6.9

1.8

@ 82.2~

19.65

73.42

35.68

@ 98.9~

13.37

37.25

20.35

Asphalthenes (wt.%)

< 0.5

1.5

< 0.5

Total sulphur (ppm)

5200

5781

1095

Total nitrogen (ppm)

3300

4858

148

Basic nitrogen (ppm)

1310

1743

141

Ni (ppm)

-

11

-

V (ppm)

0.2

16

-

Fe (ppm)

16

7

1.1

Concarbon residue (wt.%) Viscosity (cSt)

Table 3. Catalyst Characterization. A T R e-

B a s e e-

63

63

Density (g.mL-1)

0.86

0.89

Surface area (m2.g-1)

125

131

Crystallinity (%)

20

21

RE203 (wt.%)

2.94

2.28

Ni (ppm)

5066

2702

V (ppm)

5845

2028

MAT

(wt.%)

Besides determining the spent catalyst carbon content through total combustion in the LECO equipment, the coke species were separated into three

291

different fractions as follows: i)the soft coke I, obtained after Soxhlet extraction with CH2C12 after 6 h at 40~ ii)the soft coke II, soluble in CH2C12 after acid attack of the catalyst structure with HF and iii)the CH2C12 insoluble part of coke. Details of the extraction procedure adopted can be found elsewhere [6], a similar approach using CHC13 and HC1 was discussed by Koon et al [7]. 3. RESULTS AND DISCUSSION Figure 1 presents the differences in crackability of the three feedstocks. The yields of the main products are depicted in figure 2A-D. For a given CTO, the paraffinic ATR showed higher yields of the more valuable products (LPG and naphtha) and a higher bottoms conversion. Although the heavy ATR conversion levels were close to the HC~ + CGO blend, coke and gas yields were much higher.

Fig. 1. Conversion v~. CTO. (1"1)HGO + CGO, (O) ATR, (*) paraffinic ATR. The average hydrogen-to-coke ratio (H/K) calculated from the chromatographic analysis of the flue gas (Fig. 3) depends on the coke nature which is dependent on catalyst, feedstock and operational conditions, and on the amount of relative light hydrocarbons dragged to the regenerator due to lack of stripper efficiency. For a given feedstock, the stripper efficiency could be denoted based on the H/K value. The higher the CTO the higher the amount of hydrocarbons to be removed in the stripper, resulting in lower stripper efficiency and, thus a high H/K. For this reason H/K should be compared at iso-CTO.

292

The effect of feedstock could be observed comparing the results of the paraffinie ATR with the blend HGO+CGO, since the same catalyst and operational conditions were used in both cases. The paraffinic ATR presents a lower H/K compared to the blend HGO + CGO. In order to better understand this difference in H/K, coked catalyst samples were characterized in all cases, aiming to separate the impact of the amount of dragged hydrocarbons from differences in the coke nature, which depends on the coke on catalyst level as well as on the yields profiles (and type of hydrocarbons) produced during cracking. The H/K data for the heavy ATR was rather scattered.

Fig. 2. (A) Dry gas, (B) LPG, (C) naphtha, and (D) coke yields v s . CTO. (1"!)HGO + CGO, (O) heavy ATR, (*) paraffinic ATR.

293

In order to quantify the amount of entrapped molecules in the catalyst structure, selected coked samples obtained at the spent catalyst standpipe were submitted to a direct Soxhlet treatment with CH2C12. The amount of soluble coke recovered atter a direct Soxhlet treatment is higher for the parrafinic ATR (close to 10wt.%) compared to the 6wt.% direct recovered in the HGO + CGO case. The heavy ATR has only 2wt.% of compounds in this category. This result indicates that there is a relation between the chemical composition of the cracked products profile and the dessorption step during stripping. It is interesting to observe that the paraffinic ATR produces about 15 wt.% more naphtha at CTO = 8 then the other feedstocks. As the compounds in the dry gas and LPG range are easily desorbed from the catalyst structure and the heavy compounds could not be recovered, the compounds in the naphtha range are most probably the main products recovered during stripping. This explains why the total amount of strippable material is higher for the paraffinic ATR feedstock (see Fig. 2). Nevertheless, the relatively high value of soluble coke recovered after a direct Soxhlet treatment for this later feedstock indicates that there is still advantageous to improve stripper efficiency. 11"-" 10 ~ 9

4

7--

9

61,

o

s-" 4

2

~

~

" 9

i,

4

OI3

mO *i .

I

.........

.

6

:

8

.

.

10

.

;

12

CTO

Fig. 3. Hydrogen to coke ratio vs. CTO. (I-i) HGO + CGO, (4) heavy ATR, (*) paraffinic ATR. The amount of soluble coke recovered after acid treatment decreases with the total coke content deposited over pure zeolites [8]. In these experiments, within a narrower range, the same thing was observed (Fig. 4). The higher the coke content, the lower the relative amount of CH2C12 soluble species trapped in the catalyst structure.

294

For the heavy ATR feedstock, which has the highest carbon residue and asphalthenes contents more than 80 wt.% of coke is insoluble in CH2C12 after acid treatment. The insoluble coke compounds are directly correlated to the carbon residue (Fig. 5) suggesting that the heavy components which characterize the carbon residue act as insoluble coke precursors under FCC conditions. Although the insoluble coke content was the highest for the heavy ATR, the H/K was situated in the broad range of 5 to 10 and was comparable to the other feedstocks. 60

5o

"~ .~,

D

40

D 30

0

u 20 ,.Q

lO o

' ""'""'"

0

!

i

i

l

0.5

1

1.5

2

c o k e o n catalyst (wt.%)

Fig. 4. CH2C12 soluble coke after HF treatment vs'. coke on catalyst for ( D ) H G O + CGO, ( O ) heavy AT1L ( * ) paraffinic ATR. 90 8s "0

8O

"'~' 75 o

7O

9 65 J~

_= 6 0 5~ 55

[]

.E

9

50, 0

~

2

.

;

4

,

.

.

6

RCR (wt.%)

Fig. 5. Insoluble coke yield and vs. carbon residue for ( D ) H G O + CGO, ( 0 ) heavy ATR, (*) paraffinic ATR.

295

Those results indicates that the insoluble coke is more alkylated and/or naphthenic for the heavy ATR (Figs 3 and 5). The low hydrogen-to-coke ratio obtained for the paraffinic ATR compared to the HCO + CGO feedstock (Fig. 3) is in accordance with the lower soluble coke yield observed for the paraffinic ATR.

4. C O N C L U S I O N S The feedstock type has an impact in the stripper efficiency and on the coke nature. The amount of products in the naphtha range (paraffinic ATR ease) is directly related to the amount of hydrocarbons to be recovered in the stripper. The CH2Clz insoluble coke directly correlates to the amount of carbon residue. The comparison between the insoluble coke results and the average hydrogen-to-carbon ratio of coke indicates that the insoluble coke is more alkylated and/or naphthenic for the heavy ATR feedstock.

Acknowledgements The authors thanks C.A. Henriques and J.L.F. Monteiro (UFRJ/COPPE/NUCAT) for the spent catalyst analysis and the collaboration of L.C. Casaveehia, A. Jamhour, A.V. Naldoni and M.L. Galvfio (UN-SIX/PQ).

REFERENCES [1] B.E. Reynolds, J.L. Rogers and R.A. Broussard, Evaluation of resid conversion options, NPRA 1997, AM-00-97-51, 16-18th March, 1997, San Antonio, TX, USA. [2] D.L. Johnson and R.C. S~nior, FCC catalyst stripper, US Patent 5,531,884 (1996). [3] R.R. Rail and B. Demulder, New internal for maximizing performance of FCC catalyst strippers, 12th Refining Seminar, 10/Out/2000, SanFrancisco, CA, USA, 2000. [4] B.W. Hedrick, J.P. Koebel, I.B. Cetinkaya and K. Puppala, Improved catalyst stripping from cold flow modeling, Akzo-Nobel Catalyst Symposium Ecomagic, 2001, F-9. [5] E. Morgado Jr., L.F. Leite, O.C. Pravia and M.A. Torem, Hydrocarbon Engineering, January, 2002. [6] M. Guisnet and P. Magnoux, Appl. Catal. A 54 (1989) 1. [7] C.L. Koon, F. Akbar, R. Hughes, Y.R. Tyagi, M. Castro Diaz, S.C. Martin, P.J. Hall and C.E Shape, Trans. IChemE, 78, part A (2000) 738. [8] H.S. Cerqueira, P. Magnoux, D. Martin and M. Guisnet, Appl. Catal. A 208 (2000) 359.

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 02004 Elsevier B.V. All rights reserved.

297

FCC Catalyst with High LPG Yield and Lower Gasoline Olefin Content Z.-H. Qiu, Y.-B. Lu and C.-Y. Li

Research Institute of Petroleum Processing, Sinopec, Beijing 100083, China

In recent years, great attention has been paid to fluid catalytic cracking (FCC) processes capable of acting as gasoline and light olefins producers. The Research Institute of Petroleum Processing (RIPP) has carried out extensive exploratory works in this area that have resulted in the development of the maximized gas and gasoline (MGG) process aimed at maximizing liquefied petroleum gas rich in C3 and C4 olefins and high octane gasoline yields. A number of units operating in MGG mode have been put into production to meet the local requirements in China. In order to get cleaner burning, a new specification to control inferior components in motor gasoline was implemented. According to this specification, the olefins content in motor gasoline should be less than 35 v%. This limit has put tremendous pressure on refineries to reduce the olefin components in gasoline, especially for processes like MGG since reducing olefins in gasoline is in contradiction with increasing liquefied petroleum gas (LPG) olefinicity and gasoline octane. In this paper, work on improving the catalyst for the MGG process to satisfy the above requirements will be discussed. 1. INTRODUCTION In recent years, great attention has been paid to shifting the focus of the fluid catalytic cracking (FCC) process from not only acting as a gasoline producer, but also to being a major source of light olefins. The Research Institute of Petroleum Processing (RIPP) has carried out extensive exploratory works in this area and has developed new catalysts and commercialized a series of related processes [1-3]; the maximized gas and gasoline (MGG) process is one of them. This process is aimed at maximizing liquefied petroleum gas (LPG) and gasoline yields, where the gasoline has high octane and the LPG is

298

rich in C 3 and C4 olefins. The commercial trial was carried out in August 1992, and promising results were obtained. In fact, the sum of LPG + gasoline reached 78-81 wt%, while the olefin coment in LPG was 65~75 wt%. In addition, gasoline octane was high, with RON of 92-95 and MON of 81-83. Because a FCC unit can be easily converted to MGG operation with minimal cost, a number of units operating in MGG mode have been placed into production in China to meet the local requirements. However, with the growth of the vehicle market and fuel consumption, there is a need to improve gasoline quality in order to get cleaner air. Gasoline in China has passed through three stages from leaded, to unleaded, to cleaner burning. Leaded gasoline was replaced by unleaded gasoline in year 2000. This approached the first step towards improving gasoline qualities. In order to get cleaner burning, a new specification controlling inferior components in automobile gasoline was set up. According to the new specification, the olefins content in motor gasoline should be less than 35 v%. This specification has been in place in Beijing, Shanghai, and Guangzhou since July 1, 2000, and it will be extended to the whole area in 2003. The new gasoline standard has placed tremendous pressure on refineries to reduce olefin components of gasoline, especially for process like MGG, since reducing olefins in gasoline is in contradiction with increasing LPG olefinicity and gasoline octane. In this paper, work on improving the catalyst for the MGG process to satisfy the above requirements will be discussed.

2. EXPERIMENTAL 2.1. Catalyst Samples Preparation Several commercial FCCs and zeolites were used in this study" Catalyst RMG and RAG were produced by Qilu Catalyst Plant, Shandong province, China, according to CN 1072201A and CN 1085825A, they were specially formulated for the MGG process. ZRP zeolite was also produced by Qilu Catalyst Plant, China, according to CN 1058382A. Acidity modified Y zeolite samples, MY-I~3, were prepared using the method described in CN 1325940A. Cat-l~3 were prepared by mixing MY-I~3 zeolite samples and ZRP zeolite with an alumina-kaolin matrix, respectively. Based on Cat-2, a pore structure modified matrix [4 ] was involved to form Cat4. The physicochemical properties of catalyst samples, Cat-l-~ 4, are in Table 1. Before testing, all catalyst samples were steam-deactivated at 800 ~ in 100% steam for 8 hours.

2.2. Catalyst Samples Testing and Analyses The catalyst samples, Cat-l~ 3, were tested in a Fix- Fluidized- Bed (FFB) unit. The feedstock used in the test was 100% AR (paraffinic base). The

299

catalyst: oil ratios in the experiments were varied from 3 to 6 by changing the amount of feed injected. The amount of catalyst used in the experiments was 120g. The reaction temperature was 520 ~ Cat-4 was tested in an AdvancedCatalytic- Equipment (ACE) unit. The feedstock was 70wt%VGO + 30wt%VR. The amount of catalyst used in the experiments was 9g, and the reaction temperature was also 520 ~ Table 1 Main properties of catalyst samples Sample Na20, wt% Cat- 1 0.12 Cat-2 0.14 Cat-3 0.16 Cat -4 0.14

A1203, wt% 52.0 51.6 51.8 52.0

RE203, Surfacearea wt% m2/g 3.3 253 3.4 254 3.7 260 3.6 272

Porevolume, ml/g 0.19 0.18 0.17 0.21

Analyses of the different products (gas, liquid, and solid) were made with the following analysis methods. The gaseous products were analyzed in an HP6890 gas chromatograph (HP-GC) equipped with a flame ionization detector (FID). The liquid products were analyzed using a simulated distillation technique in an HP-GC equipped with an FID for hydrocarbon detection.

3. RESULTS AND DISCUSSION

Conventional FCC can produce up to 65 wt% LPG plus gasoline, in which the ratio of LPG to gasoline is about 1:5, and the overall conversion is ~ 75 wt% (based on fresh feed). For some FCCUs, where an octane promoter such as ZSM-5 additive was employed, the ratio of LPG to gasoline increased a slightly, but the overall conversion remained virtually unchanged. If the unit operates at conversion beyond 75 wt%, over-cracking might take place, producing excessive coke and dry gas. Moreover, if a high yield of olefm in LPG and a good quality gasoline are desired, thermal cracking effects should be minimized. With the RMG/RAG series catalyst the MGG technology fulfilled the above objectives successfully. A proprietary shape selective ZRP zeolite was involved in the RMG/RAG catalyst series, and the catalysts have the following characteristics" high activity, good hydrothermal stability, outstanding bottoms crackability, metals resistance and excellent olefm selectivity. In RMG/RAG catalyst series, several catalysts were designed to serve the different feedstocks, operation conditions and yields pattern. Among them RAG-1 was designed for higher gasoline yield and RAG-2 for higher LPG yield. Some of the commercial test results are listed in Table 2.

300

Table 2 Typical commercial results with RMG/RAG catalysts Lanzhou ............. Yangzhou Refinery Feedstocks VGO + 21.6wt% VR . . . . . AR (paraffin base) RMG RAG-1 RAG-2 .Catalyst .... 534 538 530 534 534 Reaction temp., ~ Product yields, wt% 4.94 5.24 4.19 H2 -~C2 3.97 5.13 C3 --~ C4 26.78 27.38 27.18 28.31 33.41 C5 -~ gasoline 48.07 51.87 47.72 49.28 40.03 LCO 12.36 4.94 11.23 6.30 13.17 Coke 8.20 9.64 8.60 10.40 8.76 Loss 0.62 1.04 0.33 0.47 0.44 Conversion, wt% 87.64 95.06 88.77 93.70 86.83 LPG + gasoline, wt% 74.85 79.25 74.90 77.59 73.44 Olefin yields, wt% C38.79 9.73 9.88 10.39 11.17 Ca7.26 7.22 10.05 10.61 10.70 ,,

,

,

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

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In order to reduce the olefin content in MGG gasoline and still have the high octane of MGG gasoline, it is necessary to obtain gasoline rich in iso-paraffins and with a certain amount of aromatics. Yet, according to the requirements of the refinery, the yield of LPG, especially propylene, should not decline significantly. The above requirements are contradictory in some way. Research on altering the MGG operation conditions and designing new catalysts were carried out recently in RIPE An important challenge in catalyst design was faced to increase the selective hydrogen transfer activity [5] and enhance the capability for producing light olefin simultaneously. It was reasoned that a more pore-opened matrix with properly adjusted acidities [4] might enhance the bottoms cracking capability of the catalyst. Also, with the use of a modified Y-zeolite, more linear olefins could probably be provided for the selective cracking to light olefins. Some of the research results are shown in Table 3-5 and Figure 1-2. Figure 1 shows acidity test results for the modified Y-zeolite samples, MY-I~3 and for a conventional Y (CY) zeolite sample. It can be seen that the total acidity, strong acidity and acid sites density are greater in the modified MY-l~3 zeolite samples than in the conventional Y zeolite. Cat-l~3 containing MY-I~3 respectively were especially formulated to serve the different feedstocks and operation conditions. In these catalysts different ZRP/Y zeolite ratios were used. As a result, the yields of LPG, gasoline and the group compositions of gasoline were changed. A commercial RAG catalyst was used as reference catalyst. It can be seen

301

from Table 3 that C a t - 1-~3 exhibit similar yield structure and conversions at the some MGG reaction conditions, as the commercial RAG catalyst. The olefin content in MGG gasoline of C a t - 1-~3 is reduced, while the gasoline octane (both RON and MON) remains unchanged due to an increase in isoparaffins and aromatics, indicating that selective hydrogen transfer reactions were enhanced by the MY zeolite (Table 4). Table 3 Yiel d structure of modified MGG catalysts (FFB results ) . . . . . . . . . . . . . . . . . . Catalyst Commercial RAG Cat- 1 Cat- 2 Feedstocks AR (paraffinic base) Reaction temp., ~ 520 Yields, wt% 2.48 2.30 H2 -~ C2 2.28 C3 --- Ca 29.14 29.34 29.73 C5 "~ gasoline 39.08 38.13 39.29 LCO 11.55 11.45 11.54 Slurry 10.25 11.50 9.55 Coke 7.69 7.10 7.60 Conversion, wt% 78.20 77.05 78.91 LPG+Gasoline, wt% 68.22 67.47 69.02 Table 4 Grou p compositions and octane number Of gasoline Catalyst Commercial RAG Group Composition, wt% Paraffins 5.72 Iso-paraffins 22.04 Olefins 30.41 Naphthenes 7.71 Aromatics 34.15 Gasoline RON 93.5 MON 81.8

. . . Cat- 1 5.17 23.95 27.88 7.81 35.18 93.8 81.6

. . . ........Cat- 2

Cat- 3

2.09 26.35 42.84 11.69 9.97 7.06 78.34 69.19

Cat- 3

4.81 24.49 25.43 7.77 37.50 94.5 82.0

4.51 28.50 26.17 7.67 33.13 94.8 81.8

Table 5 Yields 9 f light olefins Of modified MGG catalysts (FFB results) . . . . . . . . . . . . . . . Catalyst Commercial RAG Cat - 1 Cat- 2 C2=, wt% 1.01 1.18 1.14 C3 =, wt% 11.31 11.43 11.44 C4=, wt% 8.43 8.00 7.72 C2=+ C3=+ C4=, wt% 20.75 20.61 20.30 LPG olefinicity, % 71 70 68 i C4=/i C4~ 0.49 0.45 0.39

Cat - 3 0.84 9.86 7.23 17.92 68 0.39

302

Fig. 1 Results of acidity modified Y-zeolites

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Fig. 2. The catalysts hydrothermal (800~ 100%H20) stability. Table 5 shows that the enhancement of selective hydrogen transfer activity can also be seen from the decrease of i C4= to i C4~ ratio, however, the LPG olefinicities of C a t - 1~3 are only slightly lower than the reference catalyst. The above results satisfied the requirements of the refinery. Figure 2 shows that the hydrothermal stability of the new catalysts are even better than the one of the highly hydrothermally stable, commercial RAG sample. A further improved catalyst, Cat-4, was prepared and tested in an ACE unit. Test results for the catalytic performance of Cat-4 are shown in Figures 3-6. It is quite certain that, as compared with the commercial RAG catalyst at the some

303

C/O ratios, C a t - 4 exhibits improved conversion and produces more valuable products yield, especially higher propylene yield, while the gasoline olefin contents are reduced.

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304

3. S U M M A R Y Through the modification of the acidic properties of Y-zeolite and with the adoption of a more pore-opened matrix in new catalysts preparation, the requirements of reducing MGG gasoline olefins while maintaining high LPG yield can be met. Promising small-scale test results have been obtained and the commercialization of this catalyst preparation is in progress.

REFERENCES [ 1] Z.B. Chen, Z.T. Li and Y.Q. Huo, 14th World Petroleum Congress, 11 (1994). [2] Y.Q. Huo, Y.M. Wang, et al, AIChE National Spring Meeting, 14 (1993). [3] Z.B. Chen, X.X.Zhong, et al, "MIO- A Novel Process for Maximum Isoolefins Production.", 5th Congress of Chemical Engineering, paper 83k, San Diego, USA (1996). [4] Z.H. Qiu, W.H. Zhang, Z.B. Wang and C.Y. Li, CN Patent No.1388214A, (2003). [5] J.Q. Zhang, Ph.D dissertation, Research Institute of Petroleum Processing, China, (2001).

Studies in Surface Science and Catalysis, volume 149 M. Occelli (Editor) 9 Elsevier B.V. All rights reserved.

305

Innovations in producing light olefins by fluid catalytic cracking P. O'Connor," A. Hakuli, b and P. Imhof

a

aAkzo Nobel Catalysts, Stationsplein 4, 3800 AE Amersfoort, The Netherlands bAkzo Nobel Catalysts, Nieuwendammerkade 1-3, 1022 AB Amsterdam, The Netherlands

The fluid catalytic cracking (FCC) process has been evolving continuously, starting in the 1940's as the main gasoline process, and serving the additional task of converting residual feedstock's in the 1970's. Nowadays, the FCC process is employed also in the production of light olefins. Innovations in FCC process, catalysis, and synergistic combinations thereof are emerging and they enhance the capabilities of the production of light olefins. At the same time, the versatility of the process for resid conversion and/or distillate fuels production is maintained. The increasing importance of olefins production in the FCC must be addressed within the framework of future trends in specifications and compositions of transportation fuels. FCC remains an important producer of FCC gasoline and light cycle oil, and an important producer of intermediates for alkylate, MTBE, and isomerate.

1. INTRODUCTION 1.1. Propylene Demand and Supply

Several incentives are affecting the growing need for the production of propylene by fluid catalytic cracking. While the need for propylene is growing faster than that of ethylene, the co-production of propylene from steam crackers (~70% of supply) is expected to decline as plants are optimized to produce higher-value ethylene. Also, ethylene and butylenes (and possibly, also C5 isoolefins) will continue to become higher valued specialty products from the refinery and FCC process.

306

The bulk of the additional propylene will need to be produced by changing the ratios of FCC product streams. This involves optimization of the host catalyst system, and the application of small pore zeolites containing additives and catalysts. Moreover, addition of new technologies will be needed to further increase the propylene production. It has even been suggested that "on-purpose" propylene technologies may become economically viable, including dehydrogenation and the conversion of methanol to olefins [ 1]. 1.2. Shifts in transportation fuels

Changes in fuels and fuels specifications are having a great impact on the refinery. For instance, due to the growth of diesel consumption and structural over-supply of gasoline, there are several refineries in Europe, which are attempting to reduce their gasoline yield. Furthermore, the traditional quality of FCC gasoline is no longer acceptable for the gasoline pool due to the tighter sulfur and olefins specifications [2]. In Europe, "dieselization" and the drive to reduce gasoline production have induced many refiners to convert potential FCC products to lighter olefins (C3-- and C4:). In certain cases, gasoline may be re-constructed via high severity catalytic cracking to light olefins and iC4, which are fed to processes such as alkylation, isomerization and/or polymerization, yielding "green" (low sulfur and low aromatics) gasoline components. In the longer term, the aromatics content of fuels will become a very important quality constraint [2]. Also, in the case of gasoline, the trend is that the improved internal combustion engines (ICE), and certainly gasoline or naphtha based fuel cells (Gsln FC), will run more efficiently with lower aromatics fuel, and the demand for octane will become far less critical [3,4]. 1.3. FCC as a versatile process producing light Olefins and clean fuels

In this paper we address the co-evolution of process and catalysis, leading up to the present state-of-the-art in producing light olefins by FCC. Recent innovations in FCC process, catalysis, and synergistic combinations thereof are introduced. These enhance the capabilities for the production of light olefins, and maintain the versatility of the process for resid conversion and/or distillate fuels production. The latest developments in catalyst research and design are particularly emphasized. We stress that the increasing importance of the production of light olefins in the FCC can and must be achieved in harmony with the trends in future specifications and compositions of transportation fuels. Specifically for FCC, this means that the light olefins production needs to be maximized, while minimizing the sulfur and aromatics to the transportation fuel pools [2].

307

2. THE E V O L V I N G FCC PROCESS 2.1. FCC Process Developments The FCC process has been evolving continuously, starting in the 1940's as the main gasoline producer. In the 1970's, FCC started to convert residual feedstock. Nowadays, FCC is also playing an increasing role in the production of light olefins. The FCC process hardware and operation have co-evolved with the catalyst and changing economic and environmental requirements. Recent key developments in the FCC process and hardware include [5,6]" 9 9 9 9 9 9 9 9 9 9 9 9 9 9 9 9 9 9

Short contact time riser reactor Improved feed distribution and atomization Feed pre vaporization or "supercritical" injection Multiple feed injection Quick product disengaging and separation from catalyst Quick product quenching More efficient stripping Downer (down flow) reactor hnproved regenerator efficiency, lower inventory hnproved control of combustion (CO, CO2, SOx, NOx) Improved air grid designs Catalyst coolers (internal and external heat-removal) Power recovery from flue gas Improved high flux standpipes High efficiency cyclone separators Cyclones without diplegs Third-, fourth- and fifth- stage particulate capture systems Erosion and high temperature resistant metallurgy

These developments have lead to dramatic reductions in the size (elevation, volume, and/or catalyst inventory) and cost of a FCC unit. 2.2. Catalyst developments An important part of the process improvements were enabled by catalyst developments. [7,8]. In the 1950's a remarkable advancement was made when zeolites were recognized for application in catalysis [9]. In the 1960's Mobil [10] introduced rare earth stabilized faujasite zeolites into FCC catalysts leading to a very substantial increases in conversion and gasoline production (Table 1). However, the high hydrogen transfer of faujasite catalysts strongly reduced the yield of light olefins (Table 2).

308

Table 1 Improvements by zeolite cracking catalysts [ 10]. Conversion (vol%) Silica-alumina gel 56 REHX 68 REHY 75

Gasoline (vol%i 40 52 58

Table 2 Loss !n light Olefins yield by zeolite cracking catalysts. Gasoline (vol%) Conversion (vol%) Silica-alumina gel 75.5 47.5 REHY 85.5 61.0

C3= (vol%) 8.5 5.9

In the 1970's the FCC unit design and operation evolved to regain some of the lost octane and light olefin yields by higher reactor temperature operation and riser cracking, unfortunately usually at the cost of dry gas production. In the 1980's the advent of Short Contact Time (SCT) FCC, closed cyclones, and introduction of the small pore ZSM-5 zeolites resulted in improved octanes and light olefin yields, while limiting the incremental dry gas production. The function of ZSM-5 is based on shape selectivity, which means that preferentially (linear) C5+ olefins and possibly also n-alkanes or side chains are cracked to lighter olefins. In the absence of ZSM-5, C5+ olefins would undergo hydrogen transfer over the faujasite zeolite, ending up as (iso)paraffins in the gasoline fraction [ 11,12]. Increase in the reactor temperature and the addition of ZSM-5 tend to increase the rate of cracking of gasoline olefins, relative to the rate of hydrogen transfer. However, it has been observed that there is little synergy between the two methods for boosting light olefins yield. This is apparently due to a depletion of the precursor olefins (Table 3) [13]. A simplified reaction scheme about the competing reactions involved in light olefins production is shown in Figure 1. Mainly the higher (linear) olefins are the reactants, which can be converted to light olefins. However, besides cracking and isomerization, these higher linear olefins can also undergo other reactions, such as hydrogen transfer and aromatization. Therefore, it can be concluded that the maximization of light olefins in the FCC is consistent with the objective of reducing the quantity of aromatics being produced in the FCC.

309

Table 3 Operat!on at higher reactor temperatures and effect of ZSM-5 addition. Dry gas (wt%) Gasoline(wt%) Base case 2.6 45.0 Higher temperature 3.3 44.2 ZSM-5 additive 2.6 42.7

C3= (wt%) 4.2 6.3 5.2

Gasoil, Gasoline

Feed

Naphthenes

Paraffin's

. ~

Olefins & Paraffin's

Light Olefins

Fig. 1. FCC Reaction model: Production of aromatics vs. light olefins.

3. PRESENT STATE OF THE ART 3.1. Stabilization of small pore zeolite activity Since the first commercial trial [14] with a ZSM-5 additive in 1984, several further improvements in the technology have occurred [15]. A very significant innovation has been the stabilization of ZSM-5 with phosphorus. The stabilization by phosphorus has resulted in roughly a doubling of the activity of the ZSM-5 crystal, when 2-3 wt% of crystal is applied in the FCC inventory. This increased effectiveness has led to a much wider use of ZSM-5 additives in FCC. Phosphorus retards dealumination of the zeolite and a larger fraction of its acidity and activity is retained after hydrothermal deactivation (Table 4), [16]. This is despite the fact that phosphorus slightly reduces the initial acidity of the zeolite. It has also been found that the particle size of ZSM-5 influences the yields: small particles favor gasoline loss and LPG increase. This effect is more pronounced with ZSM-5 particles of higher acidity [16]. As might be expected, the benefits of phosphorus stabilization also are valid for other (small

310

pore alumino-silicate) zeolites. Phosphorus stabilization has been successfully applied to stabilize faujasite (Y-type) and MCM-41 type zeolites [17]. Table 4 Phosphorus stabilization of ZSM-5. ZSM-5 Relative activity* Fresh 100 1 steamed** 1 wt-% P, fresh 80 1 wt-% P, steamed** 6 .....

* n-hexane cracking at 811 K; ** 1033K, 5h, 100% steam, atmospheric

3.2. Optimization

of the FCC catalyst and ZSM-5

additive system

The interaction between the hydrogen transfer activity of the faujasite Y based zeolite FCC catalyst and the ZSM-5 containing additives is critical [18, 19]. The effectiveness of ZSM-5 decreases when the hydrogen transfer activity of the host FCC catalyst increases. This phenomenon is illustrated in Figure 2, in which host catalysts with high and low rare earth on zeolite (Y) was blended with ZSM-5 additive, and tested in a short contact time fluid bed test (FST). At constant conversion comparison, the yield of propylene will be maximized when a host catalyst is used with a minimum rare earth concentration and hydrogen transfer activity. Unfortunately this coincides with a higher dry gas make and a lower catalyst activity. Furthermore, we have observed that there seems to be a limit to the propylene yield that can be achieved with these catalyst systems.

................................................................................ "o o

>- 2.5

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0.5

........ , 0

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Fig. 2. Propylene yield vs. ZSM-5-to-total zeolite ratio

311

3.3. Systems with very high ZSM-5 concentration in inventory Because of the limitation in supply of higher olefins precursors to ZSM5, there is only a limited benefit from adding more ZSM-5 crystal in the FCC inventory above a certain concentration [ 18, 20]. Also, we argue that a loss in ZSM-5 effectiveness can be observed if the crystal is too highly concentrated in some parts of the catalyst system. This is illustrated in Figure 3, where the amount of ZSM-5 crystal required to obtain a constant production of propylene is given as a function of various ZSM-5 addition strategies. To summarize, we propose that two distinct issues influence the effectiveness of ZSM-5" 9 First, at the macro- and micro scale (defined here as the catalyst system and individual catalyst particle respectively), the cracking rate of ZSM-5 can be limited by the (local) supply of linear olefins. This implies that above a certain conversion the cracking rate per % crystal will deteriorate strongly. 9 Secondly, the effectiveness of the stabilization of the zeolite crystal can deteriorate when increasing the amount of crystal that is embedded in the additive or catalyst. This second effect has been addressed also by Ziebarth et al. [21 ]. Since it was found that the phosphorus needed for ZSM-5 stabilization interacts also with other active components in FCC catalyst [22], ZSM-5 was stabilized ex situ before incorporation. The ex situ stabilization is achieved by treating the zeolite crystal with a phosphate source followed by a fixation step. Subsequently, the stabilized zeolite is added to the catalyst composition, preferably just before (spray) drying.

120 100

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FCC with ZSM-5~'~

FCC + Additives