= 1 + 0.8447 x (0.913)1'2058 x (728.2)"0'0557 x (446y°M44 RI(20) =1.5105
Refractive Index (RI) @ 60°C (140°F) RI(60) = 1 + 0.8156 x (SG)L2392 x (VABP °K)'a0576 x (MW)'0'0007 RI(60) = 1 + 0.8156 x (0.913)L2392 x (728.2)"0'0576 x (446)"0'0007 Rl{60) = 1.4963
Hydrogen (H2) Content, wt% H2 = 52.825 - 14.26 x RI(20) - 21.329 x (SG) - 0.0024 x (MW) - 0.052 x (S) + 0.757 x In (v) H, = 52.825 - 14.26 x 1.5105 - (21.329 x 0.913) - (0.0024 x
FCC Feed Characterization
75
446) - (0.052 x 0.48} - (0.872 - In (7.37)) H2 = 12.23 wt%
Aromatic (CA) Content, wt% CA = -814.136 + (635.192 x RI(20)) - (129.266 x (SG)) + (0.013 x (MW)) - (0.34 x (S)) + (0.872 x In (v)) CA = -814.136 + (635.192 x 1.5105) - (129.266 x 0.913) + (0.013 x 446) - (0.34 x 0.48) + (0.872 x In 7.37) CA = 19.19 wt%
Where:
SG AP °C VABP °C VABP °K S V
= Specific gravity at 20°C = Aniline Point, °C = Volumetric Average Boiling Point, °C = Volumetric Average Boiling Point, °K = Sulfur, wt% = Viscosity at 100°C
For FCC feeds, particularly the ones containing residue, the TOTAL correlation is more accurate at predicting aromatic carbon content than he n-d-M correlation. Table 2-9 illustrates this comparison. One option s to calculate MW, RI(20)» CA, and H2 from the TOTAL correlation, and use either the n-d-M or API method to calculate the wt% naphthene CN) and wt% paraffin (Cp).
n-d-M Method
The n-d-M correlation is an ASTM (D-3238) method that uses efractive index (n), density (d), average molecular weight (MW), and ulfur (S) to estimate the percentage of total carbon distribution in the aromatic ring structure (% CA), naphthenic ring structure (CN), and paraffin chains (% Cp). Both refractive index and density are either measured or estimated at 20°C (68°F). Appendix 4 shows formulas used to calculate carbon distribution. Note that the n-d-M method alculates, for example, the percent of carbon in the aromatic ring
6
Fluid Catalytic Cracking Handbook Table 2-9 Comparison of TOTAL Correlations with Other Methods
Correlation
arbon Content (%C) n-d-M API TOTAL Hydrogen Content (%H) Linden Fein-Wilson-Winn Modified Winn TOTAL Molecular Weight (MW) API Maxwell Kesler-Lee TOTAL efractive Index (RI) API @ 20°C Lindee-Whitter @ 20°C TOTAL @ 20°C TOTAL @ 60°C
Absolute Average Deviation
Bias Maximum Deviation
5.14 2.88 0.93
4.67 2.53 0.00
12.99 9.13 3.45
0.31 0.36 0.19 0.10
-0.05 0.19 0.07 0.00
1.57 1.43 0.86 0.42
Average Deviation
62.0 63.3 61.5 10.6
-62.0 -63.6 -61.1 -0.20
180.9 175.0 176.9 44.4
0.0368
-0.0367
0.0993
0.0315 0.0021 0.0021
-0.0131 0.0 0.0
0.0303 0.0074 0.0074
ource: Dhuleaia [1]
tructure. For instance, if there was a toluene molecule in the feed, he n-d-M method predicts six aromatic carbons (86%) versus the ctual seven carbons. ASTM D-2502 is one of the most accurate methods of determining molecular weight. The method uses viscosity measurements; in the bsence of viscosity data, molecular weight can be estimated using the OTAL correlation. The n-d-M method is very sensitive to both refractive index and ensity. It calls for measurement or estimation of the feed refractive ndex at 20°C (68°F). The problem is that the majority of FCC feeds re virtually solid at 20°C and the refractometer is unable to measure
FCC Feed Characterization
77
he refractive index at this temperature. To use the n-d-M method, efractive index at 20°C needs to be estimated using published corelations. For this reason the n-d-M method is usually employed in onjunction with other correlations such as TOTAL. Example 2-3 can be used to illustrate the use of the n-d-M correlations.
Example 2-3
Using the feed property data in Example 2-1, determine MW, CA, CN nd Cp using the n-d-M method.
Step 1: Molecular weight determination by ASTM method. 1. Obtain viscosity at 100°F (37.8°C) a. Plot viscosities at 130°F (54.4°C) and 210°F (98.89°C). b. Extrapolate to 100°F, VIS = 279 SSU. 2. Convert viscosities from SUS to centistoke (csT): a. From Appendix 6, viscosity @ 100°F = 60.0 cSt. b. Viscosity @ 210°F = 7.37 cSt. 3. Obtain molecular weight: a. From Appendix 5, H = 372 and MW = 430.
Step 2: Calculate refractive index @ 20°C from the TOTAL correlation. RI(20) = 1 + 0.8447 x (SO)1'2056 x (VABP(deg C) + 273.16r°'0557 x (MW)-0'0044 RI(20) = 1 + 0.8447 x (0.913)!'2056 x 728.2^'0557 x 446^0044 RI,20) = 1.5046
Step 3: Calculate n-d-M Factors. V = 2.51 x (RI(20) - 1.4750) - (D20 - 0.8510) = 0.0271 V = 2.51 x (1.5046 - 1.4750) - (0.90 - 0.8510) = +0.0271 w = (D20 - 0.8510) - 1.11 (RI(20J - 1.4750) = +0.0226 w - (0.90 - 0.8510) - 1.11 x (1.5046 - 1.4750) = +0.0226
8
Fluid Catalytic Cracking Handbook Because V is positive: %CA = 20.16
Because w is positive: %CR =820xw-(3xS) + -
10,000
MW 10,000
%CR = 820 x 0.0226 - 3 x 0.48 +
430
The API method is a generalized method that predicts mole fraction f paraffinic, naphthenic, or aromatic compounds for an olefin-free ydrocarbon. The development of the equations is based on dividing he hydrocarbon into two molecular ranges: heavy fractions (200 < MW < 600) and light fractions (70 < MW 276x0"913) x(!276)-°-407x(0.913)-3-3333
RI = (1 + 2 x I/I- I) l/2 /
\l/2
_ 1 + 2x0.294 V R1(20) "~l 1-0.294 J RJ(20) = 1 .500
,,^ ,7. . ^ . „ SG - 0.24- 0.222 xlog(v210- 35.5) VG = Viscosity Gravity Constant = ™ 0.755 Q=
0.913 -0.24- 0.022 xlog(50- 35.5) 0.755
VG = 0.8575 XA = g + h(Ri) + i(VG)
XA = -403.8 + 265.7 xf 1.5000-^^1 + 161. Ox 0.8575 XA = 1 1 .5 rnol% XN = d + e(Ri) + f(VG)
XN = 246.4 - 367.0 x 1.5000 XN = 31.8 mol% Xp = a + b(Ri) + i(VG)
2
+ 196.3 x 0.8575
0
Fluid Catalytic Cracking Handbook Xp = 257.37 + 101.33(Ri) + 160.988(0.8575)
Xp = 56,7 mo.1%
Where:
onstants a b c d e f g h i
= = = = = = = =
+2.5737 +1.0133 -3.573 +2.464 -3.6701 +1.96312 4.0377 +2.6568 +1.60988
The findings from TOTAL, n-d-M, and API are summarized in Table -10. The comparison illustrates how sensitive the predicted feed omposition is to the refractive index @ 20°C. For instance, using the OTAL correlation, there is a 35% drop in the aromatic content in sing RI(20) = 1.5000 instead of RI(20) = 1.5105. When using these orrelations, every effort should be made to obtain accurate and onsistent values for the refractive index at 20°C. With the refractive
Table 2-10 Comparison of the Findings Among the 3 Correlations API
n-d-M
efractive Index @ 20°C Molecular Weight
1.5000
413
430
arbon Content:
Mol%
Wt%
Aromatic Naphthene Paraffin
11.5, (14.3)* 31.8, (27.9)* 56.7, (57.8)*
TOTAL
1.5105
446 Wt% t
(20.2)*,(8.8) (20.2)*,(41.1)f (59.6)*,(50.1)t
Uses Rll2i)l from n-d-M correlation to determine composition, Uses RI^0ifrorn API correlation to determine composition.
19.2, (12.5)*
FCC Feed Characterization
81
ndex at any given temperature, the RI(2o) can be calculated from the ollowing equation. Example 2-5 illustrates the use of the equation. RI (2m = RI(t) + 6.25 x (t - 20) x 10"4 t = temp, °C
Example 2-5
With the refractive index @ 78°C = 1.4810, determine the refracive index @ 20°C. RI(2()) = 1.4810 + 6.25 x (67 - 20) x 10"4 RI(20) = 1.5104
Note that the calculated RI(20) closely matches that using the TOTAL correlation.)
Pretreatment of FCC feedstock through hydroprocessing has a number of benefits including: • • • • •
Hydrodesulfurization (HDS) Hydrodenitrogenation (HDN) Hydrodemetallization (HDM) Aromatic Reduction Conradson Carbon Removal
Desulfurization of FCC feedstocks reduces the sulfur content of FCC products and SOX emissions. In the United States, road diesel sulfur can be 500 ppm (0.05 wt%). In some European countries, for example n Sweden, the sulfur of road diesel is 50 ppm or less. In California, he gasoline sulfur is required to be less than 40 ppm. The EPA's complex model uses sulfur as a controlling parameter to reduce toxic missions. With hydroprocessed FCC feeds, about 5% of feed sulfur s in the FCC gasoline. For non-hydroprocessed feeds, the FCC gasoline sulfur is typically 10% of the feed sulfur.
2
Fluid Catalytic Cracking Handbook
The nitrogen compounds in the FCC feed deactivate the FCC atalyst activity resulting in an increase in coke and dry gas. Hydroenitrogenation reduces nitrogen compounds in FCC feeds. In the egenerator, the nitrogen and the attached heterocyclic compounds add nwanted heat to the regenerator causing a low unit conversion. Hydrodemetallization reduces the amount of nickel and, to a lesser xtent, vanadium in FCC feeds. Nickel dehydrogenates feed to molecular ydrogen and aromatics. Removing these metals allows heavier gas il cut points. Polynuclear aromatics (PNA) do not react in the FCC and tend to emain in coke. Adding hydrogen to the outer ring clusters makes them more crackable and less likely to form coke on the catalyst. Hydroprocessing reduces the Conradson carbon residue of heavy ils, Conradson carbon residue becomes coke in the FCC reactor. This xcess coke must be burned in the regenerator, increasing regenerator ir requirements.
It is important to characterize FCC feeds as to their molecular tructure. Once the molecular configuration is known, kinetic models an be developed to predict product yields. The simplified correlations bove do a reasonable job of defining hydrocarbon type and distribuon in FCC feeds. Each correlation provides satisfactory results within he range for which it was developed. Whichever correlation is used, he results should be trended and compared with unit operation. A clear understanding of feed physical properties is essential to uccessful work in the areas of troubleshooting, catalyst selection, unit ptimization, and any planned revamp.
REFERENCES
1. Dhulesia, H., "New Correlations Predict FCC Feed Characterizing Parameters," Oil & Gas Journal, January 13, 1986, pp. 51-54 ASTM, "Standard Test Method for Calculation of Carbon Distribution and Structural Group Analysis of Petroleum Oils by the n-d-M Method," ASTM Standard D-3238-85, 1985. 3. Riazi, M. R., and Daubert, T. E., "Prediction of the Composition of Petroleum Fractions," Ind. Eng. Chem. Process Dev., Vol. 19, No. 2, 1982, pp. 289-294.
FCC Feed Characterization
83
4. ASTM, "Standard Test Method for Estimation of Molecular Weight (Relative Molecular Mass) of Petroleum Oils from Viscosity Measurements," ASTM Standard D-2502-92, 1992, 5. Flanders, R. L., Proceedings of the 35th Annual NPRA Q&A Session on Refining and Petrochemical Technology, Philadelphia, Pa., 1982, p. 59. 6. Wollaslon, E. G., Forsythe, W. L., and Vasalos, I. A., "Sulfur Distribution in FCC Products," Oil & Gas Journal, August 2, 1971, pp. 64-69. 7. Huling, G. P., McKinney, J. D., and Readal, T, C, "Feed-Sulfur Distribution in FCC Products," Oil & Gas Journal, May 19, 1975, pp. 73-79. 8. Campagna, R. J., Krishna, A. S., and Yanik, S. J., "Research and Development Directed at Resid Cracking," Oil and Gas Journal, October 31, 1983, pp. 129-134. 9. Davison Div., W. R. Grace & Co., "Questions Frequently Asked About Cracking Catalyst," Grace Davison Catalagram, No. 64, 1982, p. 29. 0. Andreasson, H. U. and Upson, L. L., "What Makes Octane," presented at Katalistiks' 6th Annual FCC Symposium, Munich, Germany, May 2223, 1985. 1. Van, K. B., Gevers, A., and Blum, A., "FCC Unit Monitoring and Technical Service," presented at 1986 Akzo Chemicals Symposium, Amsterdam, The Netherlands. 2. Scherzer, J., and McArthur, D. P., "Nitrogen Resistance of FCC Catalysts," presented at Katalistiks' 8th Annual FCC Symposium, Venice, Italy, 1986. 3. Dougan, T. J., Alkemade, V, Lakhampel, B., and Brock, L. T., "Advances in FCC Vanadium Tolerance," presented at NPRA Annual Meeting, San Antonio, Texas, March 20, 1994; reprinted in Grace Davison Catalagram No. 72, 1985.
CHAPTER 3
FCC Catalysts
The introduction of zeolite in commercial FCC catalysts in the early 1960s was one of the most significant advances in the history of cat racking. Zeolite catalysts provided a greater profit with little capital nvestment. Simply stated, zeolite catalysts were and still are the iggest bargain of all time for the refiner. Improvements in catalyst echnology have continued, enabling refiners to meet the demands of heir market with minimum capital investment. Compared to amorphous silica-alumina catalysts, the zeolite catalysts re more active and more selective. The higher activity and selectivity ranslate to more profitable liquid product yields and additional crackng capacity. To take full advantage of the zeolite catalyst, refiners ave revamped older units to crack more of the heavier, loweralue feedstocks. A complete discussion of FCC catalysts would fill another book. This chapter provides enough information to select the proper catalyst nd to troubleshoot the unit's operation. The key topics discussed are: • • • • • • •
Catalyst Components Catalyst Manufacturing Techniques Fresh Catalyst Properties Equilibrium Catalyst Analysis Catalyst Management Catalyst Evaluation Additives
CATALYST COMPONENTS
FCC catalysts are in the form of fine powders with an average article size in the range of 75 microns. A modern cat cracking catalyst as four major components: • Zeolite • Matrix 84
FCC Catalysts
85
* Binder • Filler
Zeolite
Zeolite, or more properly, faujasite, is the key ingredient of the FCC atalyst. It provides product selectivity and much of the catalytic ctivity. The catalyst's performance largely depends on the nature and uality of the zeolite. Understanding the zeolite structure, types, racking mechanism, and properties is essential in choosing the "right" atalyst to produce the desired yields.
Zeolite Structure
Zeolite is sometimes called molecular sieve. It has a well defined attice structure. Its basic building blocks are silica and alumina etrahedra (pyramids). Each tetrahedron (Figure 3-1) consists of a ilicon or aluminum atom at the center of the tetrahedron, with oxygen toms at the four corners. Zeolite lattices have a network of very small pores. The pore diameter f nearly all of today's FCC zeolite is approximately 8.0 angstroms (°A). These small openings, with an internal surface area of roughly 600 square
Figure 3-1.
Silicon/aluminum-oxygen tetrahedron [15].
6
Fluid Catalytic Cracking Handbook
meters per gram, do not readily admit hydrocarbon molecules that have molecular diameter greater than 8.0°A to 10°A. The elementary building block of the zeolite crystal is a unit cell. The unit cell size (UCS) is the distance between the repeating cells n the zeolite structure. One unit cell in a typical fresh Y-zeolite lattice ontains 192 framework atomic positions: 55 atoms of aluminum and 137 atoms of silicon. This corresponds to a silica (SitX,) to alumina A12O3) molal ratio (SAR) of 5. The UCS is an important parameter n characterizing the zeolite structure.
Zeolite Chemistry
As stated above, a typical zeolite consists of silicon and aluminum toms that are tetrahedrally joined by four oxygen atoms. Silicon is n a +4 oxidation state; therefore, a tetrahedron containing silicon is eutral in charge. In contrast, aluminum is in a +3 oxidation state. This ndicates that each tetrahedron containing aluminum has a net charge f -1, which must be balanced by a positive ion. Solutions containing sodium hydroxide are commonly used in ynthesizing the zeolite. The sodium serves as the positive ion to alance the negative charge of aluminum tetrahedron. This zeolite is alled soda Y or NaY. The NaY zeolite is not hydrothermally stable ecause of the high sodium content. The ammonium ion is frequently sed to displace sodium. Upon drying the zeolite, ammonia is vaporized. The resulting acid sites are both the Bronsted and Lewis types. The Bronsted acid sites can be further exchanged with rare earth material, uch as cerium and lanthanum to enhance their strengths. The zeolite ctivity comes from these acid sites.
eolite Types
Zeolites employed in the manufacture of the FCC catalyst are ynthetic versions of naturally occurring zeolites called faujasites. There are about 40 known natural zeolites and over 150 zeolites that ave been synthesized. Of this number, only a few have found commercial pplications. Table 3-1 shows properties of the major synthetic zeolites. The zeolites with applications to FCC are Type X, Type Y, and ZSM-5. Both X and Y zeolites have essentially the same crystalline tructure. The X zeolite has a lower silica-alumina ratio than the Y eolite. The X zeolite also has a lower thermal and hydrothermal
FCC Catalysts
8?
Table 3-1 Properties of Major Synthetic Zeolites
Zeoiite Type
Pore Size Dimensions (°A)
Silica-toAlumina Ratio
Zeolite A Faujasite ZSM-5
4.1 7.4 5.2 x 5.8
2-5 3-6 30-200
Mordenite
6.7 x 7.0
10-12
Applications Detergent manufacturing Catalytic cracking and hydrocracking Xylene isomerization, benzene alkylation, catalytic cracking, catalyst dewaxing, and methanol conversion. Hydro-isomerization, dewaxing
tability than the Y zeolite. Some of the earlier FCC zeolite catalysts contained X zeolite; however, virtually all of today's catalysts contain Y zeolite or variations thereof (Figure 3-2). ZSM-5 is a versatile zeolite that increases olefin yields and octane. ts application is further discussed later in this chapter. Until the late 1970s, the NaY zeolite was mostly ion exchanged with are earth components. Rare earth components, such as lanthanum and
USY Zeolite (~ 7 Al Atoms/u.c.)
nit Cell Dimension =24.25 A (SiO2/AI2O3=54)
Equilibrium REY (-23 Al Atoms/u.c.) Unit Cell Dimension = 24.39 A (SiO2/AI2O3 « 15)
Figure 3-2. Geometry of USY and REY zeolites [14].
8
Fluid Catalytic Cracking Handbook
erium, were used to replace sodium in the crystal. The rare earth lements, being trivalent, simply form "bridges" between two to three cid sites in the zeolite framework. Bridging protects acid sites from eing ejected and stabilizes the zeolite structure. Rare earth exchange dds to the zeolite activity and thermal and hydrothermal stability. The reduction of lead in motor gasoline in 1986 created the need or a higher FCC gasoline octane. Catalyst manufacturers responded y adjusting the zeolite formulations, an alteration that involved xpelling a number of aluminum atoms from the zeolite framework. The removal of aluminum increased SAR, reduced UCS, and in the rocess, lowered the sodium level of the zeolite. These changes ncreased the gasoline octane by raising its olefinicity. This aluminumeficient zeolite was called ultrastable Y, or simply USY, because of s higher stability than the conventional Y.
eolite Properties
The properties of the zeolite play a significant role in the overall erformance of the catalyst. Understanding these properties increases ur ability to predict catalyst response to changes in unit operation. rom its inception in the catalyst plant, the zeolite must retain its atalytic properties under the hostile conditions of the FCC operation. The reactor/regenerator environment can cause significant changes in hemical and structural composition of the zeolite. In the regenerator, or instance, the zeolite is subjected to thermal and hydrothermal reatments. In the reactor, it is exposed to feedstock contaminants such s vanadium and sodium. Various analytical tests determine zeolite properties. These tests upply information about the strength, type, number, and distribution f acid sites. Additional tests can also provide information about urface area and pore size distribution. The three most common arameters governing zeolite behavior are as follows: • Unit Cell Size • Rare Earth Level « Sodium Content
Unit Cell Size (UCS). The UCS is a measure of aluminum sites or he total potential acidity per unit cell. The negatively-charged aluminum toms are sources of active sites in the zeolite. Silicon atoms do not
FCC Catalysts
89
possess any activity. The UCS is related to the number of aluminum toms per cell (N Af ) by [1]: NA, + 111 x (UCS - 24.215)
The number of silicon atoms (Nsi) is; Nsi = 192 - NA,
The SAR of the zeolite can be determined either from the above two quations or from a correlation such as the one shown in Figure 3-3. The UCS is also an indicator of zeolite acidity. Because the alumium ion is larger than the silicon ion, as the UCS decreases, the acid ites become farther apart. The strength of the acid sites is determined y the extent of their isolation from the neighboring acid sites. The lose proximity of these acid sites causes destabilization of the zeolite tructure. Acid distribution of the zeolite is a fundamental factor ffecting zeolite activity and selectivity. Additionally, the UCS measurement can be used to indicate octane potential of the zeolite. A lower UCS presents fewer active sites per unit cell. The fewer acid ites are farther apart and, therefore, inhibit hydrogen transfer reactions, which in turn increase gasoline octane as well as the production of C3 and lighter components (Figure 3-4). The octane increase is due o a higher concentration of olefins in the gasoline. Zeolites with lower UCS are initially less active than the conventional rare earth exchanged zeolites (Figure 3-5). However, the ower UCS zeolites tend to retain a greater fraction of their activity under severe thermal and hydrothermal treatments, hence the name ultrastable Y. A freshly manufactured zeolite has a relatively high UCS in the ange of 24,50°A to 24.75°A. The thermal and hydrothermal environment of the regenerator extracts alumina from the zeolite structure and, herefore, reduces its UCS. The final UCS level depends on the rare arth and sodium level of the zeolite. The lower the sodium and rare arth content of the fresh zeolite, the lower UCS of the equilibrium atalyst (E-cat).
Rare Earth Level. Rare earth (RE) elements serve as a "bridge" o stabilize aluminum atoms in the zeolite structure. They prevent the
0
Fluid Catalytic Cracking Handbook
Figure 3-3.
Silica-alumina ratio versus zeolite unit cell size,
luminum atoms from separating from the zeolite lattice when the atalyst is exposed to high temperature steam in the regenerator. A fully rare-earth-exchanged zeolite equilibrates at a high UCS, whereas a non-rare-earth zeolite equilibrates at a very low UCS in the ange of 24.25 [3]. All intermediate levels of rare-earth-exchanged eolite can be produced. The rare earth increases zeolite activity and
FCC Catalysts
24.24
24.28
91
24.32
24.36
24.32
24.36
Unit Cell Size, A
6.0 5.5
>s t
1 5.0 o 4.5 «*>
4.0
24.20
24.24
24.28 Unit Cell Size, A
Figure 3-4.
Effects of unit cell size on octane and C3-gas make [4].
2
Fluid Catalytic Cracking Handbook
90 80
1520°F,20% steam in air.
0
10
20
30
40
50
60
70
80
90 100
Time, hrs
Figure 3-5. Comparison of activity retention between rare-earth-exchanged eolites versus USY zeolites. (Source: Grace Davison Octane Handbook.)
gasoline selectivity with a loss in octane (Figure 3-6). The octane loss s due to promotion of hydrogen transfer reactions. The insertion of rare arth maintains more and closer acid sites, which promotes hydrogen ransfer reactions. In addition, rare earth improves thermal and hydrohermal stability of the zeolite. To improve the activity of a USY zeolite, he catalyst suppliers frequently add some rare earth to the zeolite.
Sodium Content. The sodium on the catalyst originates either from eolite during its manufacture or from the FCC feedstock. It is important or the fresh zeolite to contain very low amounts of sodium. Sodium decreases the hydrothermal stability of the zeolite. It also eacts with the zeolite acid sites to reduce catalyst activity. In the egenerator, sodium is mobile. Sodium ions tend to neutralize the trongest acid sites. In a dealuminated zeolite, where the UCS is low 24.22°A to 24.25°A), the sodium can have an adverse affect on the gasoline octane (Figure 3-7). The loss of octane is attributed to the drop in the number of strong acid sites. FCC catalyst vendors are now able to manufacture catalysts with a odium content of less than 0.20 wt%. Sodium is commonly reported as
FCC Catalysts
5
§3
Yield of Gasoline, % ~ I
4
3
2
Gasoline octane (R+MV2
1
0 0
2
4
6
8
10
12
Rare Earth, wt% Figure 3-6. Effects of rare earth on gasoline octane and yield.
he weight percent of sodium or soda (Na2O) on the catalyst. The proper way to compare sodium is the weight fraction of sodium in the zeolite, This is because FCC catalysts have different zeolite concentrations. UCS, rare earth, and sodium are just three of the parameters that re readily available to characterize the zeolite properties. They proide valuable information about catalyst behavior in the cat cracker. f required, additional tests can be conducted to examine other eolite properties.
Matrix
The term matrix has different meanings to different people. For ome, matrix refers to components of the catalyst other than the eolite. For others, matrix is a component of the catalyst aside from he zeolite having catalytic activity. Yet for others, matrix refers to he catalyst binder. In this chapter, matrix means components of the atalyst other than zeolite and the term active matrix means the omponent of the catalyst other than zeolite heaving catalytic activity.
4
Fluid Catalytic Cracking Handbook MOTOR OCTANE VS. SODIUM OXIDE 81.5
-
,
;
0.3
0.4
0.5
81.0
O 5 80.5
80.0 0.2
0.6
Na2O, wt% on catalyst
RESEARCH OCTANE VS. SODIUM OXIDE
94
93
92
91
0
1
2
3
4
5
Na2O, wt% on zeolite
igure 3-7. Effects of soda on motor and research octanes: motor octane s. sodium oxide [11]; research octane vs. sodium oxide [4].
FCC Catalysts
95
Alumina is the source for an active matrix. Most active matrices used n FCC catalysts are amorphous. However, some of the catalyst suppliers ncorporate a form of alumina that also has a crystalline structure. Active matrix contributes significantly to the overall performance f the FCC catalyst. The zeolite pores are not suitable for cracking f large hydrocarbon molecules generally having an end point > 900°F 482°C); they are too small to allow diffusion of the large molecules o the cracking sites. An effective matrix must have a porous structure o allow diffusion of hydrocarbons into and out of the catalyst, An active matrix provides the primary cracking sites. The acid sites ocated in the catalyst matrix are not as selective as the zeolite sites, ut are able to crack larger molecules that are hindered from entering he small zeolite pores. The active matrix precracks heavy feed moleules for further cracking at the internal zeolite sites. The result is a ynergistic interaction between matrix and zeolite, in which the activity ttained by their combined effects can be greater than the sum of their ndividual effects [2]. An active matrix can also serve as a trap to catch some of the anadium and basic nitrogen. The high boiling fraction of the FCC eed usually contains metals and basic nitrogen that poison the zeolite. One of the advantages of an active matrix is that it guards the zeolite rom becoming deactivated prematurely by these impurities.
Filler and Binder
The filler is a clay incorporated into the catalyst to dilute its activity. Kaoline [Al2(OH)2, Si2O5] is the most common clay used in the CC catalyst. One FCC catalyst manufacturer uses kaoline clay as a keleton to grow the zeolite in situ. The binder serves as a glue to hold the zeolite, matrix, and filler ogether. Binder may or may not have catalytic activity. The importance f the binder becomes more prominent with catalysts that contain high oncentrations of zeolite. The functions of the filler and the binder are to provide physical ntegrity (density, attrition resistance, particle size distribution, etc.), heat transfer medium, and a fluidizing medium in which the more mportant and expensive zeolite component is incorporated. In summary, zeolite will effect activity, selectivity, and product uality. An active matrix can improve bottoms cracking and resist
6
Fluid Catalytic Cracking Handbook
vanadium and nitrogen attacks. But a matrix containing very small ores can suppress strippablity of the spent catalyst and increase ydrogen yield in the presence of nickel. Clay and binder provide hysical integrity and mechanical strength.
The manufacturing process of modern FCC catalyst is divided into wo general groups—incorporation and "in-situ" processes. All catalyst uppliers manufacture catalyst by an incorporation process that requires making zeolite and matrix independently and using a binder to hold hem together. In addition to the incorporation process, Engelhard also manufactures FCC catalyst using an "in-situ" process in which the eolite component is grown within the pre-formed miscrospheres. The ollowing sections provide a general description of zeolite synthesis.
Conventional Zeolite (KEY, REHY, HY)
NaY zeolite is produced by digesting a mixture of silica, alumina, nd caustic for several hours at a prescribed temperature until crystalization occurs (Figure 3-8). Typical sources of silica and alumina are odium silicate and sodium aluminate. Crystallization of Y-zeolite ypically takes 10 hours at about 210°F (100°C). Production of a uality zeolite requires proper control of temperature, time, and pH f the crystallization solution. NaY zeolite is separated after filtering nd water-washing of the crystalline solution. A typical NaY zeolite contains approximately 13 wt% Na2O. To nhance activity and thermal and hydrothermal stability of NaY, the odium level must be reduced. This is normally done by the ion xchanging of NaY with a medium containing rare earth cations and/ r hydrogen ions. Ammonium sulfate solutions are frequently employed s a source for hydrogen ions. At this state of the catalyst synthesis there are two approaches for urther treatment of NaY. Depending on the particular catalyst and the atalyst supplier, further treatment (rare earth exchanged) of NaY can e accomplished either before or after its incorporation into the matrix. Post-treatment of the NaY zeolite is simpler, but may reduce ion xchange efficiency.
Filtrate to waste treatment
Figure 3-8, Typical manufacturing steps to produce FCC catalyst.
Na^aoHte rystallization 00 F, 12-24 Hr
8
Fluid Catalytic Cracking Handbook
USY Zeolite
An ultrastable or a dealuminated zeolite (USY) is produced by eplacing some of the aluminum ions in the framework with silicon. The conventional technique (Figure 3-9) includes the use of a high emperature (1,300-1,500°F [704-816°C]) steam calcination of HY zeolite, (13%Na O,A,c
NAY
NHf4
2
24.68 A)
- EXCHANGES
NHY
I
(3%Na£ 90)
STEAM CALCINE 114OO DEG. F
USY
JVH.+
(3%Na O,A0* 24,50A!
. EXCHANGES
LOW-SODA USY
(
Figure 3-9. Synthesis of USY zeolite (NAY),
FCC Catalysts
9§
Acid leaching, chemical extraction, and chemical substitution are all orms of dealumination that have become popular in recent years. The main advantage of these processes over conventional dealumination s the removal of the nonframework or occluded alumina from the zeolite cage structure. A high level of occluded alumina residing in he crystal is thought to have an undesirable impact on product selectivities by yielding more light gas and LPG; however, this has not been proven commercially. In the manufacturing of USY catalyst, the zeolite, clay, and binder are slurried together. If the binder is not active, an alumina component having catalytic properties may also be added. The well-mixed slurry solution is then fed to a spray dryer. The function of a spray dryer is o form microspheres by evaporating the slurry solution, through the use of atomizers, in the presence of hot air. The type of spray dryer and the drying conditions determine the size and distribution of catalyst particles.
Engelhard Process
Engelhard's "in-situ" FCC catalyst technology is mainly based on growing zeolite within the kaolin-based particles as shown in Figure 3-9A. The aqueous solution of various kaolins is spray dried to form microspheres. The microspheres are hardened in a high-temperature (1,300°F/704°C) calcination process. The NaY zeolite is produced by digestion of the microspheres, which contain metakaolin, and mullite with caustic or sodium silicate. Simultaneously, an active matrix is ormed with the microspheres. The crystallized microspheres are iltered and washed prior to ion exchange and any final treatment.
With each shipment of fresh catalyst, the catalyst suppliers typically mail refiners an inspection report that contains data on the catalyst's physical and chemical properties. This data is valuable and should be monitored closely to ensure that the catalyst received meets the agreed specifications. A number of refiners independently analyze random samples of the fresh catalyst to confirm the reported properties. In addition, quarterly review of the fresh catalyst properties with the catalyst vendor will ensure that the control targets are being achieved.
00
Fluid Catalytic Cracking Handbook
he particle size distribution (PSD), sodium (Na), rare earth (RE), and urface area (SA) are some of the parameters in the inspection sheet hat require close attention.
article Size Distribution (PSD)
The PSD is an indicator of the fluidization properties of the catalyst, n general, fluidization improves as the fraction of the 0-40 micron articles is increased; however, a higher percentage of 0-40 micron articles will also result in greater catalyst losses. The fluidization characteristics of an FCC catalyst largely depend on he unit's mechanical configuration. The percentage of less than 40 microns in the circulating inventory is a function of cyclone efficiency. n units with good catalyst circulation, it may be economical to minimize he fraction of less than 40-micron particles. This is because after a few ycles, most of the 0-40 microns will escape the unit via the cyclones, The catalyst manufacturers control PSD of the fresh catalyst, mainly hrough the spray-drying cycle. In the spray dryer, the catalyst slurry must be effectively atomized to achieve proper distribution. As illurated in Figure 3-10, the PSD does not have a normal distribution hape. The average particle size (APS) is not actually the average size f the catalyst particles, but rather the median value.
urface Area (SA), M2/g
The reported surface area is the combined surface area of zeolite nd matrix. In zeolite manufacturing, the measurement of the zeolite urface area is one of the procedures used by catalyst suppliers to ontrol quality. The surface area is commonly determined by the mount of nitrogen adsorbed by the catalyst. The surface area correlates fairly well with the fresh catalyst activity. Upon request, catalyst suppliers can also report the zeolite surface area. his data is useful in that it is proportional to the zeolite content of he catalyst.
odium (Na), wt%
Sodium plays an intrinsic part in the manufacturing of FCC catalysts. s effects are well known and, because it deactivates the
FCC Catalysts
Figure 3-10.
101
Particle size distribution of a typical FCC catalyst.
zeolite and reduces the gasoline octane, every effort should be made o minimize the amount of sodium in the fresh catalyst. The catalyst nspection sheet expresses sodium or soda (Na2O) as the weight percent on the catalyst. When comparing different grades of catalysts, t is more practical to express the sodium content on the zeolite.
Rare earth (RE) is a generic name for 14 metallic elements of the Ianthanide series. These elements have similar chemical properties and are usually supplied as a mixture of oxides extracted from ores such as bastnaesite or monazite. Rare earth improves the catalyst activity (Figure 3-11) and hydrohermal stability. Catalysts can have a wide range of rare earth levels.
02
Fluid Catalytic Cracking Handbook
0
1
2
3
Rare Earth, wt% Figure 3-11. Effect of rare earth on catalyst activity.
epending on the refiner's objectives. Similar to sodium, the inspection heet shows rare earth or rare earth oxide (RE2O3) as the weight ercent of the catalyst. Again, when comparing different catalysts, the oncentration of RE on the zeolite should be used.
EQUILIBRIUM CATALYST ANALYSIS
Refiners send E-cat samples to catalyst manufacturers on a regular asis. As a service to the refiners, the catalyst suppliers provide nalyses of the samples in a form similar to the one shown in Figure -12. Although the absolute E-cat results may differ from one vendor o another, the results are most useful as a trend indicator. The tests performed on E-cat samples provide refiners with valuable nformation on unit conditions. The data can be used to pinpoint otential operational, mechanical, and catalyst problems because the hysical and chemical properties of the E-cat provide clues on the nvironment to which it has been exposed. The following discussion describes each test briefly and examines the ignificance of these data to the refiner. The E-cat results are divided nto catalytic properties, physical properties, and chemical analysis.
C wt% 0.23 0.23 0.16 0.23 0.22 0.20 0.24 0.15 0.24 0 0 0 0 0 0 0 0 0
1.3 1.2 1.2
Fe ppm 5600 5600 5600 5600 5600 5600 5600 5600 5600
9
1.3 1.4 1.3 1.2 1.4 1.2
G.F.
2.2 1.9 3.1 2.6 3.2 2.6 2.3 2.8 2.9
C.F.
9 0 9 8 9 9 7 0
ata
0 0 0 2 0 0 2 0 4
0-40 wt% 10 7 8 9 6 9 10 7 10
r TM
0-80 wt% 63 61 67 69 65 67 71 64 67 Cu Sb UCS RE203 ppm ppm 24.27 1.79 25 416 23 1,80 446 24 440 ' 24.27 1.79 24 r™44i 24 1.79 24.25 445 23 [_ 420 1.80 179 24 24.27 458 25 432 1.79 24 24.27 409 1.76
0-20 wt%
Figure 3-12. Typical E-cat analysis.
V Ni ppm ppm 1997 4106 4093 1948 4051 1940 4099 1974 4017 1942 3962 1910 3892 1893 3893 L_J885 1873 3875
S.A. P.V. ABD m2/gm, cc/gm gm/cc 147 0.30 0.83 148 0.83 0.28 0.84 147 0.29 148 0.83 0.29 0.28 0.83 148 0,29 0.84 150 148 0.28 0.85 148 0.85 0.29 0.28 0.84 148
130 130 130 130 130 132 131 130 130]
Z
70 72 69 68 70 69 67 71 69
APS
M
1 1 1 1 1 1 1 1 1
AI2O3 ppm 28. 29. 29. 28. 28. 28. 28. 28. 28.
04
Fluid Catalytic Cracking Handbook
Catalytic Properties
The activity, coke, and gas factors are the tests that reflect the elative catalytic behavior of the catalyst.
Conversion (Activity)
The first step in E-cat testing is to burn the carbon off the sample. he sample is then placed in a MAT unit (Figure 3-13), the heart of which is a fixed bed reactor. A certain amount of a standard gas oil eedstock is injected into the hot bed of catalyst. The activity is eported as the conversion to 430°F (221°C) material. The feedstock's uality, reactor temperature, catalyst-to-oil ratio, and space velocity are our variables affecting MAT results. Each catalyst vendor uses slightly ifferent operating variables to conduct microactivity testing, as ndicated in Table 3-2. In commercial operations, catalyst activity is affected by operating onditions, feedstock quality, and catalyst characteristics. The MAT eparates catalyst effects from feed and process changes. Feed conaminants, such as vanadium and sodium, reduce catalyst activity. -cat activity is also affected by fresh catalyst makeup rate and egenerator conditions.
Coke Factor (CF), Gas Factor (GF)
The CF and GF represent the coke- and gas-forming tendencies of an -cat compared to a standard steam-aged catalyst sample at the same onversion. The CF and GF are influenced by the type of fresh catalyst nd the level of metals deposited on the E-cat. Both the coke and gas actors can be indicative of the dehydrogenation activity of the metals on he catalyst. The addition of amorphous alumina to the catalyst will tend o increase the nonselective cracking, which forms coke and gas.
hysical Properties
The tests that reflect physical properties of the catalyst are surface rea, average bulk density, pore volume, and particle size distribution.
urface Area (SA), M2/g
For an identical fresh catalyst, the surface area of an E-cat is an ndirect measurement of its activity. The SA is the sum of zeolite and
FCC Catalysts Star dard FCCl Feed n-Q Equilibrium Catalysts
!
L
105
Syringe Pump
J -"1^3 way Valve
@
Coke Burn Off -
"I I Reactor Furnace
control
Temp
Tempcontrol
ri i— is l~W
Preheat Zone Catalyst y
rJFlow Meter L^«—}50%
This method can also be used to calculate the catalyst retention actor. The above equations assume steady-state operation, constant nit inventory, and constant addition and loss rate.
FCC Catalysts
115
Catalyst management is a very important aspect of the FCC process. election and management of the catalyst, as well as how the unit is perated, are largely responsible for achieving the desired product. roper choice of a catalyst will go a long way toward achieving a uccessful cat cracker operation. Catalyst change-out is a relatively imple process and allows a refiner to select the catalyst that maximizes he profit margin. Although catalyst change-out is physically simple, requires a lot of homework as discussed later in this section. As many catalyst formulations are available, catalyst evaluation hould be an ongoing process. However, it is not an easy task to valuate the performance of an FCC catalyst in a commercial unit ecause of continual changes in feedstocks and operating conditions, n addition to inaccuracies in measurements. Because of these limitaons, refiners sometimes switch catalyst without identifying he objectives and limitations of their cat crackers. To ensure that a roper catalyst is selected, each refiner should establish a methodology hat allows identification of "real" objectives and constraints and nsures that the choice of the catalyst is based on well-thoughtut technical and business merits. In today's market, there are over 20 different formulations of FCC catalysts. Refiners should evalute catalyst mainly to maximize profit opportunity and to minimize sk. The "right" catalyst for one refiner may not necessarily be "right" or another, A comprehensive catalyst selection methodology will have the ollowing elements: 1. Optimize unit operation with current catalyst and vendor a. Conduct test run b. Incorporate the test run results into an FCC kinetic model c. Identify opportunities for operational improvements d. Identify unit's constraints e. Optimize incumbent catalyst with vendor 2. Issue technical inquiry to catalyst vendors a. Provide test run results b. Provide E-cat sample c. Provide processing objectives d. Provide unit limitations
16
Fluid Catalytic Cracking Handbook
3, Obtain vendor responses a. Obtain catalyst recommendation b. Obtain alternate recommendation c. Obtain comparative yield projection 4, Obtain current product price projections a. For present and future four-quarters 5, Perform economic evaluations on vendor yields a. Select catalysts for MAT evaluations 6, Conduct MAT of selected list a. Perform physical and chemical analyses b. Determine steam deactivation conditions c. Deactivate incumbent fresh catalyst to match incumbent E-cat d. Use same deactivation steps for each candidate catalyst 7, Perform economic analysis of alternatives a. Estimate commercial yield from MAT evaluations 8, Request commercial proposals a. Consult at least two vendors b. Obtain references c. Check references 9, Test the selected catalysts in a pilot plant a. Calibrate the pilot plant steaming conditions using incumbent E-cat b. Deactivate the incumbent and other candidate catalysts c. Collect at least two or three data points on each by varying the catalyst-to-oil ratio 10. Evaluate pilot plant results a. Translate the pilot plant data b. Use the kinetic model to heat-balance the data c. Identify limitations and constraints 11. Make the catalyst selection a. Perform economic evaluation b. Consider intangibles-research, quality control, price, steady supply, manufacturing location c. Make recommendations 12. Post selection a. Monitoring transition-% changeover b. Post transition test run c. Confirm computer model
FCC Catalysts
117
13. Issue the final report a. Analyze benefits b. Evaluate selection methodology
There is a redundancy of flexibility in the design of FCC catalysts. Variation in the amount and type of zeolite, as well as the type of ctive matrix, provide a great deal of catalyst options that the refiner an employ to fit its needs. For smaller refiners, it may not be practical o employ pilot plant facilities to evaluate different catalysts. In this ase, the above methodology can still be used with emphasis shifted oward using the MAT data to compare the candidate catalysts. It is mportant that MAT data are properly corrected for temperature, soaking time," and catalyst strippability effects.
For many years, cat cracker operators have used additive compounds or enhancing cat cracker performance. The main benefits of these dditives (catalyst and feed additives) are to alter the FCC yields and educe the amount of pollutants emitted from the regenerator. The dditives discussed in this section are CO promoter, SOX reduction, SM-5, and antimony.
CO Promoter
The CO promoter is added to most FCC units to assist in the ombustion of CO to CO2 in the regenerator. The promoter is added o accelerate the CO combustion in the dense phase and to minimize he higher temperature excursions that occur as a result of afterburning n the dilute phase. The promoter allows uniform burning of coke, articularly if there is uneven distribution between spent catalyst and ombustion air. Regenerators operating in full or partial combustion can utilize the enefits of the CO promoter. The addition of the promoter tends to ncrease the regenerator temperature and NOx emission. The metallurgy f the regenerator internals should be checked for tolerance of the igher temperature. The active ingredients of the promoter are typically the platinum roup metals. The platinum, in the concentration of 300 ppm to 800
18
Fluid Catalytic Cracking Handbook
pm, is typically dispersed on a support. The effectiveness of the romoter largely depends on its activity and stability. Promoter is frequently added to the regenerator two to three times day, normally at a rate of 3 to 5 pounds (1 to 2.3 kg) promoter per on of fresh catalyst. The concentration of platinum required in a unit nventory is about 0.5 to 1.5 ppm. The promoter addition rate may e increased if antimony solution is being used to passivate the nickel. The use of CO promoter, particularly during unit start-ups, improves he stability of the regeneration operation. However, not every cat racker can justify combustion-promoted operation. Heat balance, vailability of combustion air, NOX emission metallurgical limits, and he presence of CO boiler are some of the factors that should be onsidered before using combustion promoter. For example, in units perating with low oxygen levels and partial combustion, a promoted ystem could increase carbon on regenerated catalyst (CRC). This is ecause CO combustion reaction competes with carbon burning reacion for the available oxygen. The combustion of CO to CO2 will also ncrease NOX emissions. This is largely due to the oxidation of intermediates such as ammonia and cyanide gases into nitrogen oxide (NO).
SOX Additive
The coke on the spent catalyst entering the regenerator contains ulfur. In the regenerator, the sulfur in the coke is converted to SO2 nd SO3. The mixture of SO2 and SO3 is commonly referred to as SOX, nd approximately 80% to 90% of SOX is SO2, with the rest being SO3, The SOX leaves the regenerator with the flue gas and is eventually ischarged to the atmosphere. Coke yield, thiophenic sulfur content f the feed, the regenerator operating condition, and the type of FCC atalyst are the major factors affecting SOX emissions. The environmental impact of SOX emissions has gained much ttention over the past ten years. The United States Environmental rotection Agency (EPA) New Source Performance Standards (NSPS) went into effect in 1989. The ruling covers new, modified, and recontructed FCC units since January 1994. It should be noted that the outhern California Air Quality Management District (SCAQMD) oard has established a limit of 60 kilograms of SOX per 1,000 barrels f feed for the existing FCC units.
FCC Catalysts
119
There are three common methods for SOX abatement. These are flue as scrubbing, feedstock desulfurization, and SOX additive. The SOX dditive is often the least costly alternative, which is the approach racticed by many refiners. The SOX additive, usually a metal oxide, is added directly to the atalyst inventory. The additive works by adsorbing and chemically onding with SO3 in the regenerator. This stable sulfate species is arried with the circulating catalyst to the riser, where it is reduced r "regenerated" by hydrogen or water to yield H2S and metal oxide. able 3-3 shows the postulated chemistry of SOX reduction by a OX agent. To achieve the highest efficiency of SOX additive, it is imporant that: * Excess oxygen be available; oxygen promotes the SO2 to SO3 reaction. SOX additive will only form a metal sulfate from SOV * The regenerator temperature be lower; lower temperature favors SO2 + 1/2 O2 -> SO3 * The capturing agent be physically compatible with the FCC catalyst and easily regenerated in the riser and stripper. * CO promoter be used, which oxidizes SO2 to SO3. * There be a uniform distribution of air and spent catalyst. Air/ catalyst mixing in the regenerator can significantly affect the SOX pick-up efficiency.
Table 3-3 Mechanism of Catalytic SOX Reduction
. In the Regenerator Sulfur in Coke (S) + O2 SO2 + l/2 O2 MXO + SO,
—» —> —>
MXSO4
—> -> —>
MXS + 4H2O MXO + H2S + 3 H2O MXO + H2O
SO2 + SO3 SO3
. In the Reactor and Stripper MXSO4 + 4H2
Mxso4 + 4H2 MXS + H2O
ource: Thiel [9]
20
Fluid Catalytic Cracking Handbook
* Operation of the reactor stripper be efficient. The stripper efficiency is very important to allow the release of sulfate and the formation of H2S.
Since most of the regenerators operating in full combustion mode sually operate with 1% to 3% excess oxygen, the capturing efficiency f SOX additive is often greater in full combustion than in partial ombustion units.
ZSM-5
ZSM-5 is Mobil Oil's proprietary shape-selective zeolite that has a ifferent pore structure from that of Y-zeolite. The pore size of ZSMis smaller than that of Y-zeolite (5.1°A to 5.6°A versus 8°A to 9°A), n addition, the pore arrangement of ZSM-5 is different from Y-zeolite, s shown in Figure 3-16. The shape selectivity of ZSM-5 allows
Figure 3-16.
Comparison of Y faujasite and ZSM-5 zeolites [13].
FCC Catalysts
121
referential cracking of long-chain, low-octane normal paraffins, as well as some olefins, in the gasoline fraction. ZSM-5 additive is added to the unit to boost gasoline octane and o increase light olefin yields. ZSM-5 accomplishes this by upgrading ow-octane components in the gasoline boiling range (C7 to C1O) into ght olefins (C3, C4, C5), as well as isomerizing low-octane linear lefins to high-octane branched olefins, ZSM-5 inhibits paraffin ydrogenation by cracking the C7+ olefins. ZSM-5's effectiveness depends on several variables. The cat crackers hat process highly paraffinic feedstock and have lower base octane ill receive the greatest benefits of using ZSM-5. ZSM-5 will have ttle effect on improving gasoline octane in units that process naphthenic eedstock or operate at a high conversion level. When using ZSM-5, there is almost an even trade-off between FCC asoline volume and LPG yield. For a one-number increase in the esearch octane of FCC gasoline, there is a 1 vol% to 1.5 vol% ecrease in the gasoline and almost a corresponding increase in the PG, This again depends on feed quality, operating parameters, and ase octane. The decision to add ZSM-5 depends on the objectives and conraints of the unit. ZSM-5 application will increase load on the wet as compressor, FCC gas plant, and other downstream units. Most efiners who add ZSM-5 do it on a seasonal basis, again depending n their octane need and unit limitations. The concentration of the ZSM-5 additive should be greater than 1 % f the catalyst inventory to see a noticeable increase in the octane. An octane boost of one research octane number (RON) will typically equire a 2% to 5% ZSM-5 additive in the inventory. It should be oted that the proper way of quoting percentage should be by SM-5 concentration rather than the total additive because the activity nd attrition rate can vary from one supplier to another. There are new enerations of ZSM-5 additives that have nearly twice the activity of he earlier additives. In summary, ZSM-5 provides the refiner the flexibility to increase asoline octane and light olefins. With the introduction of reformulated asoline, ZSM-5 could play an important role in producing isoutylene, used as the feedstock for production of methyl tertiary butyl ther (MTBE).
22
Fluid Catalytic Cracking Handbook
Metal Passivation
As discussed in Chapter 2, nickel, vanadium, and sodium are the metal compounds usually present in the FCC feedstock. These metals eposit on the catalyst, thus poisoning the catalyst active sites. Some f the options available to refiners for reducing the effect of metals n catalyst activity are as follows: • • • • • •
Increasing the fresh catalyst makeup rate Using outside E-cat Employing metal passivators Incorporating metal trap into the FCC catalyst Using demetalizing technology to remove the metals from the catalyst The MagnaCat separation process (demetalizing technology), which allows discarding the "older" catalyst particles containing higher metal levels
Metal passivation in general, and antimony in particular, are discussed n the following section. In recent years, several methods have been patented for chemical assivation of nickel and vanadium. Only some of the tin compounds ave had limited commercial success in passivating vanadium. Although n has been used by some refiners, it has not been proven or as widely ccepted as antimony. In the case of nickel, antimony-based comounds have been most effective in reducing the detrimental effects f nickel poisoning. It should be noted that, although the existing ntimony-based technology is the most effective method of reducing he deleterious effects of nickel, the antimony is fugitive and can be onsidered hazardous. In this case, a bismuth-based passivator may be better choice.
ntimony
Antimony-based passivation was introduced by Phillips Petroleum n 1976 to passivate nickel compounds in the FCC feed. Antimony is njected into the fresh feed, usually with the help of a carrier such as ght cycle oil. If there are feed preheaters in the unit, antimony should e injected downstream of the preheater to avoid thermal decomposion of the antimony solution in the heater tubes. The effects of antimony passivation are usually immediate. By orming an alloy with nickel, the dehydrogenation reactions that are
FCC Catalysts
123
aused by nickel are often reduced by 40% to 60%. This is evidenced y a sharp decline in dry gas and hydrogen yield. Nickel passivation is generally economically attractive when the ickel content of the E-cat is greater than 1,000 ppm. The Phillips etroleum secondary antimony patent position is due to expire in late 1999, At that time, antimony passivation can become economically ttractive at a lower nickel level than 1,000 ppm. The antimony solution should be added in proportion to the amount f nickel present in the feed. The optimum dosage normally coresponds to an antimony-to-nickel ratio of 0.3 to 0.5 on the E-cat. Antimony's retention efficiency on the catalyst is in the range of 75% o 85% without the recycling of slurry oil to the riser. If slurry recycle s being practiced, the retention efficiency is usually greater than 90%. Any antimony not deposited on the circulating catalyst ends up in the ecanted oil and the catalyst fines from the regenerator. It is often a ood practice to discontinue antimony injection about one month prior o a scheduled unit shutdown to make sure the exposure to catalyst ust containing antimony is reduced to a minimum when wearing a alf-faced respirator.
SUMMARY
The introduction of zeolite into the FCC catalyst in the early 1960s was one of the most significant developments in the field of cat racking. The zeolite greatly improved selectivity of the catalyst, esulting in higher gasoline yields and indirectly allowing refiners to rocess more feed to the unit. With the introduction of reformulated asoline, new formulations in FCC catalyst will again help refiners meet new requirements in gasoline quality. Since there are over 120 different FCC catalyst formulations in the market today, it is important that the refinery personnel involved in at cracker operations have some fundamental understanding of catalyst echnology. This knowledge is useful in areas such as proper troublehooting and customizing a catalyst that would match the refiner's eeds. The additive technology will be expanding in coming years. he need to produce reformulated gasoline will increase demand for he shape-selective zeolite, such as ZSM-5. The pressure from environmental agencies to reduce SOX and NOX will further increase the emand for additives that reduce emissions.
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Fluid Catalytic Cracking Handbook
REFERENCES 1. Breck, D. W., Zeolite Molecular Sieves: Structure, Chemistry, and Use, New York: Wiley Interscience, 1974. 2. Hayward, C. M. and Winkler, W. S., "FCC: Matrix/Zeolite," Hydrocarbon Processing, February 1990, pp. 55-56. 3. Upson, L. L., "What FCC Catalyst Tests Show," Hydrocarbon Processing November 1981, pp. 253-258. 4. Pine, L. A., Maher, P. J., and Wachter, W. A., "Prediction of Cracking Catalyst Behavior by a Zeolite Unit Cell Size Model," Journal of Catalysis, No. 85, 1984, pp. 466-476. 5. Magnusson, J. and Pudas, R., "Activity and Product Distribution Characteristics of the Currently Used FCC Catalyst Systems," presented at Katalistiks' 6th Annual FCC Symposium, Munich, Germany, May 2223, 1985. 6. John G. S. and Mikovsky, R. J., "Calculation of the Average Activity of Cracking Catalysts," Chemical Engineering Science, Vol. 15, 1961, pp. 172-175. 7. Gaughan, J. R., "Effect of Catalyst Retention on Inventory Replacement," Oil & Gas Journal, December 26, 1983, pp. 141-145. 8. Tamborski, G. A., Magnabosco, L. M., Powell, J. W., and Yoo, J. S., "Catalyst Technology Improvements Make SOX Emissions Control Affordable," presented at Katalistiks' 6th Annual FCC Symposium, Munich, Germany, May 22-23, 1985. 9. Thiel, P. G., Blazek, J. J., "Additive R," Grace Davison Catalagram, No. 71, 1985. 10. Engelhard Corporation, "Reduced Unit Cell Size Catalysts Offer Improved Octane for FCC Gasoline," The Catalyst Report, TI-762. 11. Engelhard Corporation, "Increasing Motor Octane by Catalytic Means Part 2," The Catalyst Report, EC6100P. 12. Engelhard Corporation, "The Chemistry of FCC Coke Formation," The Catalyst Report, Vol. 7, Issue 2. 13. Majon, R. J. and Spielman, J., "Increasing Gasoline Octane and Light Olefin Yields with ZSM-5," The Catalyst Report, Vol. 5, Issue 5, 1990. 14. Davison Div., W.R. Grace & Co., Grace Davison Catalagram, No. 72, 1985. 15. Humphries, Adrian P., "Zeolite Fundamentals and Synthesis," Akzo Chemicals, 1987. 16. Davison Octane Handbook. 17. G. Yaluris and A. W. Peters, "Studying the Chemistry of the FCCU Regenerator in the Laboratory Under Realistic Conditions," Grace Davison, Columbia, MD, 1998.
CHAPTER 4
Chemistry of FCC Reactions
A complex series of reactions (Table 4-1) take place when a large as-oil molecule comes in contact with a 1,200°F to 1,400°F (650°C o 760°C) FCC catalyst. The distribution of products depends on many actors, including the nature and strength of the catalyst acid sites. lthough most cracking in the FCC is catalytic, thermal cracking eactions also occur. Thermal cracking is caused by factors such as on-ideal mixing in the riser and poor separation of cracked products the reactor. The purpose of this chapter is to: • Provide a general discussion of the chemistry of cracking (both thermal and catalytic). • Highlight the role of the catalyst, and in particular, the influence of zeolites. • Explain how cracking reactions affect the unit's heat balance.
Whether thermal or catalytic, cracking of a hydrocarbon means the reaking of a carbon to carbon bond. But catalytic and thermal crackg proceed via different routes. A clear understanding of the different echanisms involved is beneficial in areas such as: • Selecting the "right" catalyst for a given operation • Troubleshooting unit operation • Developing a new catalyst formulation Topics discussed in this chapter are: • Thermal cracking • Catalytic cracking • Thermodynamic aspects
125
26
Fluid Catalytic Cracking Handbook Table 4-1 Important Reactions Occurring in FCC
. Cracking: Paraffins cracked to olefins and smaller paraffins Olefins cracked to smaller olefins
C9Hl8 -> C4H8 + C5H10
Aromatic side-chain scission
ArC10H21
Naphthenes (cyclo-paraffins) cracked to olefins and smaller ring compounds
Cyclo-C1oH20 -> C6H12
CSH1
ArC5H9
C4H8
. Isomerization: Olefin bond shift
1-C4H8 -^ trans-2-C4H8
Normal olefins to iso-olefin
n-C5H10 —> iso-C5H10
Normal paraffins to iso-paraffin
n-C4H10
Cyclo-hexane to cyclo-pentane
C6H12 + C5H9CH3
. Hydrogen Transfer:
iso-C4Hlo
Naphthene + Olefin -» Aromatic + Paraffin
Cyclo-aromatization
. Trans-alkylation/Alkyl-group Transfer
2C6H5CH3
C6H6
. Cyclization of Olefins to Naphthenes
C7H14 -^ CH3-cyclo-C6H ii
. Dehydrogenation
n-C8H18
. Dealkylation
H6 lso-C3H 7-C6H5 -» C6H6 -\. C C3 H
. Condensation
Ar-C3H == CH2 + R,CH = CHR2 v AT Ar + 2H
~^
C
8H16
+ H
2 3
7
6
jT\I
THERMAL CRACKING
Before the advent of the catalytic cracking process, thermal cracking was the primary process available to convert low-value feedstocks into ghter products. Refiners still use thermal processes, such as delayed oking and visibreaking, for cracking of residual hydrocarbons.
Chemistry of FCC Reactions
127
Thermal cracking is a function of temperature and time. The reaction ccurs when hydrocarbons in the absence of a catalyst are exposed to high emperatures in the range of 800°F to 1,200°F (425°C to 650°C). The initial step in the chemistry of thermal cracking is the formation f free radicals. They are formed upon splitting the C-C bond. A free adical is an uncharged molecule with an unpaired electron. The upturing produces two uncharged species that share a pair of elecrons. Equation 4-1 shows formation of a free radical when a paraffin molecule is thermally cracked.
R2 V "D IV,,
H
f*
7 IV
| H
I
*/"*
V-
T
t>
V- '
.IV
' H
(4-0
H
Free radicals are extremely reactive and short-lived. They can ndergo alpha scission, beta scission, and polymerization. (Alphacission is a break one carbon away from the free radical; betacission, two carbons away.) Beta-scission produces an olefin (ethylene) and a primary free adical (Equation 4-2), which has two fewer carbon atoms [1]: J\
""""' V.'..il.^ —~ VvlT'} "•"*— V-, —
Lisy ""' "
°~~~
Vx
S~M.-y T IT'iV--- ~™ V_-ilo
\iT"1"jiii. )
The newly formed primary free radical can further undergo betacission to yield more ethylene. Alpha-scission is not favored thermodynamicaily but does occur. Alpha-scission produces a methyl radical, which can extract a ydrogen atom from a neutral hydrocarbon molecule. The hydrogen xtraction produces methane and a secondary or tertiary free radical Equation 4-3).
-» CH4 + R-CH2-CH2-CH2-CH2-'CH-CH2-CH3
(4-3)
This radical can undergo beta-scission. The products will be an lpha-olefin and a primary free radical (Equation 4-4).
28
Fluid Catalytic Cracking Handbook
R-CH2-CH2-CH2-CH2-'CH-CH2-CH3 -» R-CH2-CH2-'CH2 + H2C=CH-CH2-CH3
(4-4)
Similar to the methyl radical, the R-*CH2 radical can also extract a hydrogen atom from another paraffin to form a secondary free radical nd a smaller paraffin (Equation 4-5). R,-'CH 2 + R-CH2-CH2-CH2-CH2-CH2-CH2-CH3 -> R,-CH 3 + R-CH2-CH2-CH2-CH2-CH2-*CH-CH3
R-*CH? is more stable than H3*C. Consequently, the hydrogen extracion rate of R-*CH2 is lower than that of the methyl radical. This sequence of reactions forms a product rich in C} and G,, nd a fair amount of alpha-olefins. Free radicals undergo little branchng (isomerization). One of the drawbacks of thermal cracking in an FCC is that a high percentage of the olefins formed during intermediate reactions polymerize and condense directly to coke. The product distribution from thermal cracking is different rom catalytic cracking, as shown in Table 4-2. The shift in product istribution confirms the fact that these two processes proceed via ifferent mechanisms,
CATALYTIC CRACKING Catalytic reactions can be classified into two broad categories: * Primary cracking of the gas oil molecules • Secondary rearrangement and re-cracking of cracked products
Before discussing mechanisms of the reactions, it is appropriate to eview FCC catalyst development and examine its cracking properties. An in-depth discussion of FCC catalyst was presented in Chapter 3.
FCC Catalyst Development
The first commercial fluidized cracking catalyst was acid-treated atural clay. Later, synthetic silica-alumina materials containing 10 to
Chemistry of FCC Reactions
129
Table 4-2 Comparison of Products of Thermal and Catalytic Cracking
Hydrocarbon Type
-Paraffms
Thermal Cracking
Catalytic Cracking
C2 is major product, with C3 to C6 is major product; much C3 and C3, and C4 to few n-olefins above C4; C16 olefins; little branching much branching
Olefins
Slow double-bond shifts and little skeletal isomerization; H-transfer is minor and nonselective for tertiary olefins; only small amounts of aromatics formed from aliphatics at 932°F (500°C)
Rapid double-bond shifts, extensive skeletal isomerization, H-transfer is major and selective for tertiary olefins; large amounts of aromatics formed from aliphatics at 932°F (500°O
Naphthenes
Crack at slower rate than paraffins
If structural groups are equivalent, crack at about the same rate as paraffins
Alkyl-aromatics
Cracked within side chain
Crack next to ring
ource: Venuto [2]
15 percent alumina replaced the natural clay catalysts. The synthetic ilica-alumina catalysts were more stable and yielded superior products. In the mid-1950s, alumina-silica catalysts, containing 25 percent lumina, came into use because of their higher stability. These synthetic atalysts were amorphous; their structure consisted of a random array f silica and alumina, tetrahedrally connected. Some minor improvements in yields and selectivity were achieved by switching to catalysts uch as magnesia-silica and alumina-zirconia-silica.
mpact of Zeolites
The breakthrough in FCC catalyst was the use of X and Y zeolites uring the early 1960s. The addition of these zeolites substantially ncreased catalyst activity and selectivity. Product distribution with a eolite-containing catalyst is different from the distribution with an morphous silica-alumina catalyst (Table 4-3). In addition, zeolites are 1,000 times more active than the amorphous silica alumina catalysts.
130
Fluid Catalytic Cracking Handbook Table 4-3 Comparison of Yield Structure for Fluid Catalytic Cracking of Waxy Gas Oil over Commercial Equilibrium Zeolite and Amorphous Catalysts
Yields, at 80 vol% Conversion
Hydrogen, wt% C1's + C2's, wt%
Amorphous, High Alumina
Zeolite, XZ-25
Change from Amorphous
0.08 3.8
0.04 2.1
-0.04 -1.7
Propylene, vol% Propane, vol% Total C3's
16.1 1.5 17.6
11.8 1.3 13.1
-4.3 -0.02 -4.5
Butenes, vol% -Butane, vol% -Butane, vol% Total C4's
12.2 7.9 0,7 20.8
7.8 7.2 0.4 15.4
-4.4 -0.7 -0.3 -5.4
C5-390 at 90% ASTM asoline, vol%
55.5
62.0
+6.5
Light Fuel Oil, vol% Heavy Fuel Oil, vol% Coke, wt%
4.2 15.8 5.6
6.1 13.9 4.1
+1.9 -1.9 -1.5
Gasoline Octane No.
94
89.8
-4.2
ource: Venuto [2]
The higher activity comes from greater strength and organization of he active sites in the zeolites. Zeolites are crystalline alumina-silicates having a regular pore tructure. Their basic building blocks are silica and alumina tetrahedra. Each tetrahedron consists of silicon or aluminum atoms at the center of he tetrahedron with oxygen atoms at the corners. Because silicon and luminum are in a +4 and 4-3 oxidation state, respectively, a net charge f -1 must be balanced by a cation to maintain electrical neutrality. The cations that replace the sodium ions determine the catalyst's ctivity and selectivity. Zeolites are synthesized in an alkaline environment such as sodium hydroxide, producing a soda-Y zeolite. These oda-Y zeolites have little stability but the sodium can be easily
Chemistry of FCC Reactions
131
exchanged. Ion exchanging sodium with cations, such as hydrogen or rare earth ions, enhances acidity and stability. The most widely used rare earth compounds are lanthanum (La3*) and cerium (Ce3+). The catalyst acid sites are both Bronsted and Lewis type. The catalyst can have either strong or weak Bronsted sites; or, strong or weak Lewis sites. A Bronsted-type acid is a substance capable of donating a proton. Hydrochloric and sulfuric acids are typical Bronsted acids. A Lewis-type acid is a substance that accepts a pair of electrons. Lewis acids may not have hydrogen in them but they are still acids. Aluminum chloride is the classic example of a Lewis acid. Dissolved in water, it will react with hydroxyl, causing a drop in solution pH. Catalyst acid properties depend on several parameters, including method of preparation, dehydration temperature, silica-to-alumina ratio, and the ratio of Bronsted to Lewis acid sites,
Mechanism of Catalytic Cracking Reactions
When feed contacts the regenerated catalyst, the feed vaporizes. Then positive-charged atoms called carbocations are formed. Carbocation is a generic term for a positive-charged carbon ion. Carbocations can be either carbonium or carbenium ions. A carbonium ion, CH5+, is formed by adding a hydrogen ion (H+) to a paraffin molecule (Equation 4-6), This is accomplished via direct attack of a proton from the catalyst Bronsted site. The resulting molecule will have a positive charge with 5 bonds to it. R — CH2 — CH2 — CH2 — CH3 + H+ (proton attack) -» R — C+H — CH2 — CH2 — CH3 + H2
(4-6)
The carbonium ion's charge is not stable and the acid sites on the catalyst are not strong enough to form many carbonium ions. Nearly all the cat cracking chemistry is carbenium ion chemistry. A carbenium ion, R-CH2+, comes either from adding a positive charge to an olefin or from removing a hydrogen and two electrons from a paraffin (Equations 4-7 and 4-8). R — CH. = CH — CH2 — CH2 — CH3 + H* (a proton @ Bronsted site) —>_^ jp^ ~™_™. ^^ |"j
« v-'Jcin ——• v^JrJ'-j"""
v^in.'-) —"""— v^-ii-t
\£|.— / j
32
Fluid Catalytic Cracking Handbook
R — CH2 -— CH2 — CH2 — CH3 (removal of H~ @ Lewis site) _» R _ c+H — CH2 — CH2 — CH3
(4-8}
Both the Bronsted and Lewis acid sites on the catalyst generate arbenium ions. The Bronsted site donates a proton to an olefin molecule and the Lewis site removes electrons from a paraffin moleule. In commercial units, olefins come in with the feed or are prouced through thermal cracking reactions. The stability of carbocations depends on the nature of alkyl groups tached to the positive charge. The relative stability of carbenium ions as follows [2] with tertiary ions being the most stable: Tertiary .
C ""~ V.-C+
'"•"""" V--
> P
V_^
Secondary P
\*s
P+
V~"
P
V-"'
>
Primary R
JLX.
P
V-'
> Ethyl > Methyl P+
V_--
P
V--
P+
*—•
P* V,,'
c
One of the benefits of catalytic cracking is that the primary and econdary ions tend to rearrange to form a tertiary ion (a carbon with hree other carbon bonds attached). As will be discussed later, the ncreased stability of tertiary ions accounts for the high degree of ranching associated with cat cracking. Once formed, carbenium ions can form a number of different eactions. The nature and strength of the catalyst acid sites influence he extent to which each of these reactions occur. The three dominant eactions of carbenium ions are: * The cracking of a carbon-carbon bond * Isomerization * Hydrogen transfer
Cracking Reactions
Cracking, or beta-scission, is a key feature of ionic cracking. Betacission is the splitting of the C-C bond two carbons away from the ositive-charge carbon atom. Beta-scission is preferred because the nergy required to break this bond is lower than that needed to break he adjacent C-C bond, the alpha bond. In addition, short-chain hydroarbons are less reactive than long-chain hydrocarbons. The rate of
Chemistry of FCC Reactions
133
he cracking reactions decreases with decreasing chain length. With short chains, it is not possible to form stable carbenium ions. The initial products of beta-scission are an olefin and a new carbenium on (Equation 4-9). The newly-formed carbenium ion will then continue a series of chain reactions. Small ions (four-carbon or five-carbon) can ransfer the positive charge to a big molecule, and the big molecule can crack. Cracking does not eliminate the positive charge; it stays until two ions collide. The smaller ions are more stable and will not crack, They survive until they transfer their charge to a big molecule, R _ " V,' ri+"H 11 IV. "
CH V*' 1. !••->
PH V--' 1 !••) — PH V--J. .I')
"
CH \.~^ I to
-* CH3 — CH = CH2 + C+H2 — CH2 — CH2R
(4-9)
Because beta-scission is mono-molecular and cracking is endohermic, the cracking rate is favored by high temperatures and is not equilibrium-limited.
somerization Reactions
Isomerization reactions occur frequently in catalytic cracking, and nfrequently in thermal cracking. In both, breaking of a bond is via beta-scission. However, in catalytic cracking, carbocations tend to earrange to form tertiary ions. Tertiary ions are more stable than secondary and primary ions; they shift around and crack to produce branched molecules (Equation 4-10). (In thermal cracking, free radicals yield normal or straight chain compounds.)
CH3 — CH, -— C+H — CH, — CH2R -» CH3 — C+ — CH — CH2R H
CR
or CH — CH2 — CH2R
(4-10) Some of the advantages of isomerization are:
34
Fluid Catalytic Cracking Handbook
* Higher octane in the gasoline fraction. Isoparaffins in the gasoline boiling range have higher octane than normal paraffins. * Higher-value chemical and oxygenate feedstocks in the C3/C4 fraction. Isobutylene and isoamylene are used for the production of methyl tertiary butyl ether (MTBE) and tertiary amyl methyl ether (TAME). MTBE and TAME can be blended into the gasoline to reduce auto emissions. * Lower cloud point in the diesel fuel. Isoparaffins in the light cycle oil boiling range improve the cloud point.
Hydrogen Transfer Reactions
Hydrogen transfer is more correctly called hydride transfer. It is a imolecular reaction in which one reactant is an olefin. Two examples re the reaction of two olefins and the reaction of an olefin and naphthene. In the reaction of two olefins, both olefins must be adsorbed on ctive sites that are close together. One of these olefins becomes a araffin and the other becomes a cyclo-olefin as hydrogen is moved om one to the other. Cyclo-olefin is now hydrogen transferred with nother olefin to yield a paraffin and a cyclodi-olefin. Cyclodi-olefin will then rearrange to form an aromatic. The chain ends because romatics are extremely stable. Hydrogen transfer of olefins converts hem to paraffins and aromatics (Equation 4-11). 4 CrlH2n -» 3 Cn H2n+2 + CnH2n^ olefins
—> paraffins
+ aromatic
(4-11)
In the reaction of naphthenes with olefins, naphthenic compounds re hydrogen donors. They can react with olefins to produce paraffins nd aromatics (Equation 4-12). 3 C n H 2n + CmH2m
-» 3 Cn H 2n+2
olefins
—> paraffins
+ naphthene
+ Cm H2m^6 + aromatic
(4-12)
A rare-earth-exchanged zeolite increases hydrogen transfer reactions. n simple terms, rare earth forms bridges between two to three acid ites in the catalyst framework. In doing so, the rare earth protects
Chemistry of FCC Reactions
135
hose acid sites. Because hydrogen transfer needs adjacent acid sites, bridging these sites with rare earth promotes hydrogen transfer reactions. Hydrogen transfer reactions usually increase gasoline yield and stability. The reactivity of the gasoline is reduced because hydrogen ransfer produces fewer olefins. Olefins are the reactive species in gasoline for secondary reactions. Therefore, hydrogen transfer reactions indirectly reduce "overcracking'1 of the gasoline. Some of the drawbacks of hydrogen transfer reactions are: • * • *
Lower gasoline octane Lower light olefin in the LPG Higher aromatics in the gasoline and LCO Lower olefin in the front end of gasoline
Other Reactions
Cracking, isomerization, and hydrogen transfer reactions account for he majority of cat cracking reactions. Other reactions play an imporant role in unit operation. Two prominent reactions are dehydrogenation and coking.
Dehydrogenation. Under ideal conditions (i.e., a "clean" feedstock and a catalyst with no metals), cat cracking does not yield any appreciable amount of molecular hydrogen. Therefore, dehydrogenation eactions will proceed only if the catalyst is contaminated with metals uch as nickel and vanadium.
Coking. Cat cracking yields a residue called coke. The chemistry of coke formation is complex and not very well understood. Similar o hydrogen transfer reactions, catalytic coke is a "bimolecular" eaction. It proceeds via carbenium ions or free radicals. In theory, coke yield should increase as the hydrogen transfer rate is increased. t is postulated [4] that reactions producing unsaturates and multi-ring aromatics are the principal coke-forming compounds. Unsaturates such as olefins, diolefins, and multi-ring polycyclic olefins are very reactive and can polymerize to form coke. For a given catalyst and feedstock, catalytic coke yield is a direct unction of conversion. However, an optimum riser temperature will minimize coke yield. For a typical cat cracker, this temperature is
36
Fluid Catalytic Cracking Handbook
bout 950°F (510°C). Consider two riser temperatures, 850°F and ,050°F (454°C and 566°C), at the extreme limits of operation. At 50°F, a large amount of coke is formed because the carbenium ions o not desorb at this low temperature. At 1,050°F (566°C), a large mount of coke is formed, largely due to olefin polymerization. The minimum coking temperature is within this range.
THERMODYNAMIC ASPECTS
As stated earlier, catalytic cracking involves a series of simultaneous eactions. Some of these reactions are endothermic and some are xothermic. Each reaction has a heat of reaction associated with it Table 4-4). The overall heat of reaction refers to the net or combined eat of reaction. Although there are a number of exothermic reactions, he net reaction is still endothermic. The regenerated catalyst supplies enough energy to heat the feed o the riser outlet temperature, to heat the combustion air to the flue as temperature, to provide the endothermic heat of reaction, and to ompensate for any heat losses to atmosphere. The source of this nergy is the burning of coke produced from the reaction. It is apparent that the type and magnitude of these reactions have an mpact on the heat balance of the unit. For example, a catalyst with less ydrogen transfer characteristics will cause the net heat of reaction to be more endothennic. Consequently this will require a higher catalyst circuation and, possibly, a higher coke yield to maintain the heat balance.
UMMARY
Although cat cracking reactions are predominantly catalytic, some onselective thermal cracking reactions do take place. The two proesses proceed via different chemistry. The distribution of products learly confirms that both reactions take place, but that catalytic eactions predominate. The introduction of zeolites into the FCC catalyst in the early 1960s rastically improved the performance of the cat cracker reaction roducts. The catalyst acid sites, their nature, and strength have a major influence on the reaction chemistry. Catalytic cracking proceeds mainly via carbenium ion intermediates. he three dominant reactions are cracking, isomerization, and hydrogen
n-C10H22 -> n-C7H16 + C3H6 1~C8H16 -> 2C4Hg 4C6H12 -» 3C6H14 + C6H6 cyclo-C6Hl2 + 3 1-C5H!0 -> 3n-C5H12 + C6H6 1-C4H8 -» trans-2-C4H8 n-C6H10 -» iso-C4H10 o-C6H4(CH3)2 -> m-C6H4(CH3)2 cyclo-C6H12 -» CH3-cyclo-C5H9 C6H6 + m-C6H4(CH3)2 -> 2C6H5CH3 1-C7H14 -» CH3-cyclo-C6H11 iso-C3H7-C6H5 -> C6H6 + C3H6 n-C6H14 ^ 1-C6H12 + H2 3C2H4 —> 1-C6H12 1-C4H8 + iso-C4H10 -> iso-C8H18
Specific Reaction 850°F
2.46 2.10 11.09 10.35 0.25 -0.23 0.30 1.09 0.65 1.54 0.88 -1.52 — —
950°F
— 1.05 — -1.2 3.3
980°F
— 1.10 0.65
— 0.09 -0.36
— 2.23 —
2.04 1.68 12.44 11.22 0.32 -0.20 0.33 1.00 0.65 2.11 0.41 -2.21 — —
Log KE (equilibrium constant)
H
Table 4-4 Some Thermodynamic Data for Idealized Reactions of Importance in Catalytic Crac
n Class
g
n transfer
ation
ylation ion ation genation ization Alkylation
enuto [2]
38
Fluid Catalytic Cracking Handbook
ransfer. Finally, the type and degree of reactions occurring will nfluence the unit heat balance.
REFERENCES*
. Gates, B. C., Katzer, J. R., and Schuit, G. G., Chemistry of Catalytic Processes. New York: McGraw-Hill, 1979. . Venuto, P. B. and Habib, E. T., Fluid Catalytic Cracking with Zeolite Catalysts. New York: Marcel Dekker, Inc., 1979, . Broekhoven, E. V. and Wijngaards, H., "Investigation of the Acid Site Distribution of FCC Catalysts with Ortho-xylene as a Model Compound," 1988 Akzo Chemicals FCC Symposium, Amsterdam, The Netherlands, . Koerroer, G. and Deeba, M., "The Chemistry of FCC Coke Formation," Engelhard Corporation, The Catalyst Report, Vol. 7, Issue 2, 1991.
he author also expresses appreciation to Messrs. Terry Reid of Akzo Nobel and Tom abib of Davison Div., W. R. Grace & Co., for their many helpful comments.
CHAPTER 5
Unit Monitoring and Control
The only proper way to monitor the performance of a cat cracker s by periodic material and heat balance surveys on the unit. By arrying out these tests frequently, one can collect, trend, and evaluate he unit operating data. Additionally, meaningful technical service to ptimize the unit operation should be based on regular test runs. Understanding the operation of a cat cracker also requires in-depth nowledge of the unit's heat balance. Any changes to feedstock quality, perating conditions, catalyst, or mechanical configuration will impact he heat balance. Heat balance is an important tool in predicting and valuating the changes that will affect the quantity and the quality of CC products. Finally, before the unit can produce one barrel of product, it must irculate catalyst smoothly. One must be familiar with the dynamics f pressure balance and key process controls. The main topics discussed in this chapter are: • • • •
Material Balance Heat Balance Pressure Balance Process Control Instrumentation
n the material and heat balance sections, the discussions include: • Two methods for performing test runs • Some practical steps for carrying out a successful test run • A step-by-step method for performing a material and heat balance survey • An actual case study
139
40
Fluid Catalytic Cracking Handbook
n the pressure balance section, the significance of the pressure balance n debottlenecking the unit is discussed. Finally, fundamentals of both basic" and "advanced" process controls are presented. This chapter presents the entire procedure for performing heat and weight balances. The last section of the chapter discusses the use of he distributed control system and computer in automating the process,
MATERIAL BALANCE
Complete data collection should be carried out weekly. Since changes n the unit are continuous, regular surveys permit distinction among he effects of feedstock, catalyst, and operating conditions. An accurate ssessment of a cat cracker operation requires reliable plant data. A easonable weight balance should have a 98% to 102% closure. In any weight balance exercise, the first step is to identify the input nd output streams. This is usually done by drawing an envelope(s) round the input and output streams. Two examples of such envelopes re shown in Figure 5-1. One of the key pieces of data is the composition of products leaving he reactor. The reactor effluent vapor entering the main fractionator ontains hydrocarbons, steam, and inert gases. By weight, the hydroarbons in the reactor overhead stream are equal to the fresh feed plus ecycle minus the portion of the feed that has been converted to coke. the feed can contain water, it should be analyzed for and corrected. The sources of steam in the reactor vapor are: lift steam to the andpipe, atomization steam to the feed nozzles, dome steam, and ripping steam. Some units may have other streams and the feed may ontain water. Depending on the reactor pressure, approximately 25% o 50% of the stripping steam is entrained with the spent catalyst owing to the regenerator, which should be deducted. Inert gases such as nitrogen and carbon dioxide enter the riser ntrained with the regenerated catalyst. The quantity of these inert asses is directly related to catalyst circulation rate. These gases flow hrough the gas plant and leave the unit with the off-gas from the ponge oil absorber column. They are not significant for the weight alance, but they are usually the only source of inerts in the off-gas nd should be deducted. FCC products are commonly reported, on an inert-free basis, as the olume and weight fractions of the fresh feed. In a rigorous weight
Unit Monitoring and Control
141
External Streams^-"
Figure 5-1. FCC unit input/output streams.
balance, gasoline and light cycle oil (LCO) yields and unit converion are reported based on fixed end points. The common end points are 430°F (221 °C) TBP for gasoline and 700°F TBP for LCO, Other popular cut points are 430°F (221°C) ASTM D-86 for gasoline and 650°F (343°C) or 670°F (354°C) ASTM D-86 for LCO. Using fixed
42
Fluid Catalytic Cracking Handbook
ut points isolates the reactor system from the distillation sysem performance. Conversion is defined as the volume or weight percent of feedstock onverted to gasoline and other lighter products, including coke. However, conversion is typically calculated by subtracting the volume ercent or weight percent of liquid products heavier than gasoline from resh feed, and dividing by the volume or weight of fresh feed. This s shown as follows: ~ , Feed - (light cycle oil + heavy cycle oil + decanted oil) . ,,„ Conversion m% = ^-^ —x 100
Depending on seasonal demands, the gasoline end point can range rom 380°F to 450°F (193°C to 232°C). Undercutting of gasoline ncreases the LCO product and can appear as low conversion. Thereore, it is necessary to distinguish between the apparent and true onversion. The apparent conversion is calculated before the gasoline nd point adjustment is made, and the true conversion is calculated fter the adjustment.
Testing Methods
The material balance around the riser requires the reactor effluent omposition. Two techniques are used to obtain this composition. Both echniques require that the coke yield be calculated. The first technique is to draw an envelope with the reactor effluent s the inlet stream and the product flows as the outlet streams. Streams rom other units must be included. The flow rates and composions of the entering and leaving streams are then totaled. The net is he reactor effluent. This is the method practiced by most refiners. The second technique involves direct sampling of the reactor effluent Figure 5-2). In this technique, a sample of reactor effluent is collected n an aluminized polyester bag for separation and analysis. There are several advantages and disadvantages to reactor effluent ampling;
dvantages of Reaction Mix Sampling • Allows data gathering on different sets of conditions without waiting for the recovery side to equilibrate.
lop container
Sample probe
Figure 5-2,
Reaction mix sampling [2].
Gate and ball valves
44
Fluid Catalytic Cracking Handbook
* Eliminates concern about rate and compositions of extraneous streams entering the gas plant because they are not included in the overall balance. * Eliminates concern about correcting for end points because the effluent sample is cut at the desired TBP end point. * Eliminates concern about obtaining a 100% weight balance.
Disadvantages of Reaction Mix Sampling * Possible leaks during sampling. * Possible inaccurate measurement of volume of gas and weight of liquid. * Requires qualified individuals to perform the test. » Requires separate lab to perform analyses. * Can require special procedures and be expensive.
Recommended Procedures for Conducting a Test Run
A successful test run requires a clear definition of objectives, careful lanning, and proper interpretation of the results. The following steps an be used as a guide to ensure a smooth and successful test run,
rior to the Test Run 1. Issue a memo to the involved departments: operations, laboratory, maintenance, and oil movement. Communicate the purpose, duration, and scope of the test run. Include a list of samples and the required analyses (Table 5-1). 2, Inform the units feeding the FCC. The composition of FCC feedstock should remain relatively constant during the test run. Flow meters should be zeroed and calibrated. Sample taps should be checked, particularly those that are not used regularly. 5, The sample bombs used to collect gas and LPG products should be purged, marked, and ready.
ata Collection 1. The duration of a test run is usually 8 to 12 hours. 2. Operating parameters should be specified. It should be documented which constraints (i.e., blower, wet gas compressor, etc.) the unit is operating against.
Unit Monitoring and Control
145
Table 5-1 Typical Laboratory Analysis of FCC Streams
Tests °API D-86 D-1160
Gas Oil
Sulfur
Viscosity
Metals
/
/
GC
/
/
/
Slurry Recycle
/
/
/
Decanted Oil Product
/
/
/
LCO Product
/
/
/
Gasoline Product
/
/
/
/
/
/
Feedstock /
LPG
C.,\s and C4's
Tail Gas
/
3. The sample taps must be bled adequately before samples are collected. A reliable flue gas analysis is important; an extra sample can be collected. The laboratory should retain the unused samples until all analyses are verified. 4. Pertinent operating data must be collected. A form similar to the one shown in Table 5-2 can be used to gather the data.
Mass Balance Calculations 1. The orifice plate meter factor should be adjusted for actual operating parameters. For liquid streams, the flow meters should be adjusted for °API gravity, temperature, and viscosity. For gas streams, the flow rate should be adjusted for the operating temperature, pressure, and molecular weight. 2. Chromatographs of each stream must be normalized to 100%. The GC of the off-gas must include accurate analysis of hydrogen, 3. The coke yield should be calculated using air rate and flue gas composition.
46
Fluid Catalytic Cracking Handbook Table 5-2 Operating Data
eed and Product Rates resh Feed Rate oker Off Gas CC Tail Gas PG asoline CD O
Other Pertinent Flow Rates ispersion Steam eactor Stripping Steam eactor Dome Steam ir to Regenerator
50,000 bpd (331 nrVhr) 3,000,000 scfd (3,540 m3/hr) 16,000,000 scfd (18,878 m3/hr) 11,565 bpd (77 mVhr) 30,000 bpd (199 m3/hr) 10,000 bpd (66 m3/hr) 3,000 bpd (20 mVhr) 9,000 Ib/hr (4,082 kg/hr) 13,000 Ib/hr (5,897 kg/hr) 1,200 Ib/hr (544 kg/hr) 90,000 scf/min (152,912 m3/hr)
emperature,°F/°C iser Inlet iser Outlet lower Discharge egen. Dense Phase egen. Flue Gas mbient
594/312 972/522 374/190 1,309/709 1,330/721 80/27
ressure, psig/Kp lower Discharge egen. Dome eactor Dome egenerated Catalyst Slide Valve, AP pent Catalyst Slide Valve, AP
43/296 34/234 33/227 5.8/40 6.0/41
lue Gas Analysis, Mol% O, O, O O2 2 + AT Miscellaneous Data elative Humidity resh Catalyst Makeup -Cat MAT
1.5 15.4 0.0
500 ppm -> 0.05 mol% 83.05 80% 4 tons/day 68%
Unit Monitoring and Control
147
4. The flow rate of each stream should be converted to weight units. 5. The quantity of inert gases and extraneous streams should be subtracted from the FCC gas plant products. 6. The raw mass balance should be reported, including the error, Then the feed/products should be normalized to 100%. The error will be distributed in proportion to flow rates or a known inaccurate meter will be adjusted. 7. Gasoline and LCO rates will be adjusted to standard cut points. 8. The feed characterization correlations discussed in Chapter 2 should be used to determine the composition of fresh feed.
Analysis of Results 1. The yields and quality of the desired products should be reported and compared with the unit targets. 2. The results of this test run should be compared with the results of previous test runs; any significant changes in the yields and/ or operating parameters should be highlighted. 3. The final step is to perform simple economics of the unit operation and make recommendations that improve short- and longterm unit operation.
The following case study demonstrates a step-by-step approach to performing a comprehensive material and heat balance.
A test run is conducted to evaluate the performance of a 50,000 bpd 331 m3/hr) FCC unit. The feed to the unit is gas oil from the vacuum nit. No recycle stream is processed; however, the off-gas from the elayed coker is sent to the gas recovery section. Products from he unit are fuel gas, LPG, gasoline, LCO, and decanted oil (DO). Tables 5-2 and 5-3 contain stream flow rates, operating data, and aboratory analyses. The meter factors have been adjusted for actual perating conditions. The mass balance is performed as follows: 1. Identification of the input and output streams used in the overall mass balance equation. 2. Calculation of the coke yield.
48
Fluid Catalytic Cracking Handbook Table 5-3 Feed and Product Inspections
Feed
API Gravity ulfur, Wt% Analine Point, °F/ °C RI @ 67°C Viscosity, SSU @ 150°F (65.5°C) @ 210°F (98.9°C) Distillation, °F Vol% 10 30 50 70 90 EP
Component
H, CH4 C, 2= C3
,=
C4 NC4 C4 C5+ H2S N2
O2
otal p. Gravity
Decanted Oil
Gasoline
LCO
58.5
21.5
2.4
D-86
D-86
D-1160
125 160 213 285 369 433
477 514 547 576 627 666
646 687 720 771 846 1,055
25.2 0.5 208/97.8 1.4854 109 54 D-1160 682 766 835 901 1,001 1,060
Mole% Composition of FCC Gas Plant Streams
FCC Tail Gas
15.5 35.8 17.1 11.0 1.6 4.7 0.7 0.2 1.3 1.0 2.1 7.2 1.8 100.0 0.78
LPG
17.9 31.3 16.1 10.9 23.8
100.0 0.55
FCC Gasoline
0.4 2.0 4.4 93.2
100.0
Coker Off-Gas
8.0 47.2 14.9 2.5 8.4 4.4 0.9 3.2 3.4 4.9 2.0 0.2 100.0 0.96
Unit Monitoring and Control
3, 4, 5, 6,
149
Conversion of the flow rates to weight units (e.g., Ib/hr). Normalization of the data to obtain a 100% weight balance. Determination of the component yields. Adjustment of the gasoline, LCO, and decanted oil yields to standard cut points.
nput and Output Streams in the Overall Mass Balance
As shown in Envelope 1 of Figure 5-1, the input hydrocarbon streams are fresh feed and coker off-gas. The output streams are FCC ail gas (minus inerts), LPG, gasoline, LCO, DO, and coke.
Coke Yield Calculations
As discussed in Chapter 1, a portion of the feed is converted to coke n the reactor. This coke is carried into the regenerator with the spent catalyst. The combustion of the coke produces H2O, CO, CO2, SO2, and traces of NOx. To determine coke yield, the amount of dry air to he regenerator and the analysis of flue gas are needed. It is essential o have an accurate analysis of the flue gas. The hydrogen content of coke relates to the amount of hydrocarbon vapors carried over with he spent catalyst into the regenerator, and is an indication of the eactor-stripper performance. Example 5-1 shows a step-by-step calulation of the coke yield. Example 5-1 Determination of the Unit's Coke Yield
Given: Wet air = 90,000 SCFM, Relative Humidity = 80%, Ambient Temperature = 80°F (26.7°C)
Figure 5-3 can be used to obtain percent dry air as a function of ambient emperature and relative humidity. For this example, the percentage of dry ir is 97.1% or: A- = Ami 90,000SCF x Imole x 60 Min = ,,,,., ,. • nDry Air 0.971 x — 13,817 moles/hr Min 379.5 SCF 1 hr
Flue gas rate (dry basis) is calculated from the dry air rate using nitrogen nd argon as tie elements.
50
Fluid Catalytic Cracking Handbook
„ ,, , . N (13,817 moles/hrx 0.7901) i a i _ , „ * Flue gas rate (dry = 13,145 moles/hr J basis)= 0.8305
.7901 and 0.8305 are concentrations of (nitrogen + argon) in atmospheric ry air and flue gas (from analysis), respectively.
he flow rates of each component in the flue gas stream are: « * * *
O2 out = 0.015 x 13,145 moles/hr = 197 moles/hr CO2 out = 0.154 x 13,145 moles/hr = 2,024 moles/hr SOj out = 0.0005 x 13,145 moles/hr = 7 moles/hr (N2~ + Ar) out = 0.8305 x 13,145 moles/hr = 10,917 moles/hr
n oxygen balance can be used to calculate water formed by the combuson of coke: * O2 out = 197 + 2,024 +7 = 2,228 moles/hr * 1)3 in = 0.2095 x 13,817 moles/hr = 2,895 moles/hr * O2 used for combustion of hydrogen = 2,895 - 2,228 = 667 raoles/hr
ince for each mole of O2, two moles of water are formed, the amount of ater is: * H2O formed = 667 x 2 = 1,334 moles/hr
omponents of coke are carbon, hydrogen, and sulfur. Their rates are calcuted as follows: * * « *
Carbon = 2,024 moles/hr x 12 Ibs/mole = 24,288 Ibs/hr Hydrogen = 1,334 moles/hr x 2.02 Ibs/mole = 2,695 Ibs/hr Sulfur = 7 moles/hr x 32.1 Ibs/moles = 225 Ibs/hr Coke = 24,288 + 2,695 + 225 = 27,208 Ibs/hr
* H-, content of coke, wt% = —: — x 100 = 9.9 27,231 Ibs/hr The hydrogen content of coke indicates the amount of hydrocarbon vapors arried through the stripper with the spent catalyst.
Conversion to Unit of Weight, Ibs/hr
The next step is to convert the flow rate of each stream in the verall mass balance equation to the unit of weight (e.g., Ibs/hr). xample 5-2 shows these conversions for gas and liquid streams.
Figure 5-3,
50
80
100
Dry air versus relative humidity and temperature.
Temperature ,Deg F
70
Dry Air versus Relative Humidity & Temperature
1tO
52
Fluid Catalytic Cracking Handbook
Example 5-2 Conversion of Input and Output Streams to the Unit of Weight (Ib/hr)
„ , „ . 50,000bbl 1day 141.5 350.3 Ib • Fresh Feed = —-—-- x - - x • -- x - —— day 24 hr (131.5 + 25.2) bbl = 658,964 lb/hr »
„,
3,000,000 SCF A.V 1day 1mole 27.8 lbs ni-—«___Q« 1|«,„,,. , •* . *„._„,„_ _^ V _. ,. , . . . . u _.-_...-... Sf\ x1U/m/nr A. v ._.....,.._L..._, JvJ.O 111 day 24 hr 379.5 SCF 1mole
f^r\lrpir gas CTQG — v^UJVCI —*"_!_ " !_ ........
.™ .. 16,000,000 SCF Iday Imole • FCC tail gas = —-- -x - ^x day 24 hr 379.5 SCF
22.6 Ibs x Imole
= 39,701 Ib/hr
he amount of inerts in the FCC tail gas is: 16,000,OOOSCF 1day _ _ _ „ Imole KT •N x - i- x 0.072 x 2 = —-- -day 24 hr 379.5 SCF „ 16,000,OOOSCF A n - t 1day 1mole . CO2 = —!- -x 0.021 x - ^-x - day 24 hr 379.5 SCF
281bs 3,542lb/hr x= 3,542 Ib/hr Imole 441bs , , » « , , „ x= 1,623 Ib/hr Imole
• Inert-free FCC tail gas = 39,701 - (3,542 + 1,623) = 34,537 Ib/hr .
LpG= H.565bbl x lday x
day
24 hr
141.5 X35031b = (131.5 + 123.5) bbl
„ r 30,000bbl day 141.5 350.31b • Gasoline = —-- x --x x day 24 hr (131.5 + 58.5) bbl = 326, 102 Ib/hr
day
24hr
141.5 (131.5 + 21.5)
f
bbl
3,000bbl 1day 141.5 350.31b ., «--,.,, = —x - i-x x -- = 46,2731b/hr day 24 hr (131.5 + 2.4) bbl
Unit Monitoring and Control
153
Normalization of the Data
Because a preliminary weight balance seldom has a 100% closure, t is necessary to normalize the yield to obtain a 100% weight balance, Example 5-3 shows the preliminary overall weight balance.
Example 5-3 Preliminary Overall Weight Balance
nput = Fresh Feed + Coker Off-Gas Output = FCC tail gas + LPG + Gasoline + LCO + DO + Coke • Input = 658,814 + 9,182 = 667,996 Ib/hr • Output = 34,617 + 93,656 + 326,124 + 134,973 + 46,270 + 27,231 = 662,871 lb/hr « Difference = 667,996 - 662,871 = 5,125 lb/hr
Error in mass balance = 0.8 wt%
The products are adjusted upward in proportion to theilr rates to obtain a 100% weight balance. The normalized rates: • • • • « •
Tail gas LPG Gasoline LCO DO Coke
= 34,883 Ib/hr = 94,460 lb/hr = 328,766 Ib/hr = 136,054 lb/hr = 46,626 Ib/hr = 27,440 Ib/hr
= = = =
11,658 bpd 30,230 bpd 10,077 bpd 3,023 bpd
Component Yield
The reactor yield is then determined by performing a component balance. The amount of C5+ in the gasoline boiling range is calculated by subtracting the C4 and lighter components from the total gas plant products. Example 5-4 shows the step-by-step calculation of the component yields. The summary of the results, normalized but unadusted for the cut points is shown in Table 5-4.
54
Fluid Catalytic Cracking Handbook Example 5-4 Calculation of Individual Components 0.155 x 16 MMSCFDx 2.02 ~~" 379.5x24
0.08 x 3 MMSCFDx 2.02 379,5x24
=
2
_„ 0.358x16 MMSCFDx 16 0.472x3.0 MMSCFDx 16 - C M 1 U r t CH 44 = --= 7,585 Ib/hr 379.5x24 379.5x24 C 2 ~~~ = 0-171 xl6MMSCFDx30 379.5x24 /„ _ 0.11x16 MMSCFDx 28 2
379.5x24
0.149 x 3 MMSCFDx 30 379.5x24 0.025 x 3 MMSCFD x 28 379.5x24
0.016x!6MMSCFDx44 | 0.179x 11,65]8BPDx 175.3 + ~~ 379.5x24 " 24 =
3
0.084x3 MMSCFDx 44 = 15,262 lb/hr 379.5 x 24 ^_ 0.047x16 MMSCFDx 42 0.313x11,658 BPDx 181.8 • Cr3 = + 379.5 x 24 24 0.044x3 MMSCFDx 42 O A C A / I I U / , = 30,504 Ib/hr 379.5 x 24 '
4
^0.002x16MMSCFDx58 "~ 379.5x24
| +
0.109x11,658BPDx204.6 24
0.02 x 30,230 x 204.6 MMSCFD x 42 0.032 x 3 MMSCFD x 58 24 379.5x24 =15,579 Ib/hr 0.007x16MMSCFDx58 '~ 379.5x24 =
4
[ +
0.161x11658BPDx 197.2 24
0.004x30,230x204.6 BPDx 197.2 24 = 16,95 8 Ib/hr
0.009x3 MMSCFDx 58 379.5x24
Unit Monitoring and Control
155
0.013xl6MMSCFDx56 0.238x1 l,658BPDx213.4 + 379.5x24 24 0.044x30,230x213.4 0.034x3MMSCFDx56 = 37,150 Ib/hr 24 379.5x24
Table 5-4
Normalized FCC Weight Balance Summary with Coker Gas Subtracted
Stream
Fresh Feed
bpd
ib/hr
Vol% of Feed
Wt% of Feed
50,000
658,814
100.00
100.00
Products
H7 C,
C,
c; Total C2 and lighter
H2S
C3
C?
C4 NC4
c;
Gasoline (Cs+)
LCO
DO Coke Total Apparent Conversion nerts
0.07 1.15 1.15 0.79 3.16
497 7,585 7,549 5,187 20,818
2,090 4,027 2,064 1,827 4,178
1,032 15,262 30,504 16,958 15,579 37,150
4.18 8.05 4.13 3.65 8.36
0.16 2.32 4.63 2.57 2.36 5.64
28,650 10,077
311,437 136,008
57.30 20.15
47.27 20.64
3,023
46,626 27,440 658,814
6.05
7.08 4.17 100.00 72.28
55,936
5,143
111.87 73.8
56
Fluid Catalytic Cracking Handbook
djustment of Gasoline and LCO Cut Points
As discussed earlier in this chapter, gasoline and LCO yields are enerally corrected to a constant boiling range basis. The most commonly used bases are 430°F TBP gasoline and 640°F TBP LCO end oints. Since TBP distillations are not routinely performed, they are ften estimated from the D-86 distillation data. The adjustments to the nd points involve the following: * Adding to the raw LCO all the 430°F+ in the raw gasoline and subtracting the 430°F in the LCO stream. « Adding to the raw LCO all the 650°F~ in the raw decanted oil and subtracting the 650°F~ in the decant oil stream. * Adding to the raw gasoline all the 430°F~ in the raw LCO and subtracting the 430°F* in the gasoline stream. • Adding to the raw decanted oil all the 650°F+ in the raw LCO and subtracting the 650°F~ in the decant oil stream.
Table 5-5 illustrates steps used to convert ASTM D-86 data to TBP. he laboratory usually converts D-1160 and reports the data as D-86, xtrapolation of the TBP data indicates the following: « « « •
The The The The 514
430°F+ content of the FCCU gasoline is 3 vol%, or 859 bpd. gasoline (430°F~) content of LCO is 8 vol%, or 806 bpd. 650°F+ content ofLCO is 12 vol%, or 1,209 bpd. LCO (650°F~) content of the decanted oil is 17 vol%, or bpd.
herefore, the adjusted rates are as follows: Gasoline (C5+ to 430°F TBP end point) = 28,650 - 859 + 806 = 28,597 bpd LCO (430°F to 650°F TBP end point) = 10,077 + 514 - 1,209 - 806 + 859 = 9,435 bpd DO (650°F+) = 3,023 + 1,209 - 514 = 3,718 bpd
able 5-6 shows the normalized FCC weight balance with the adjusted ut points.
Unit Monitoring and Control
Table 5-5 Conversion of ASTM Distillation to TBP Distillation for Gasoline, LCO, and Decanted Oil Gasoline TBP (From Appendix 9, TBP 50% point = 213°F) Given D-86 50% 30% 10% 70% 90% EP -
- 30% = - 10% = - IBP = - 50% = - 70% = 90% =
From Appendix 10 53°F 35°F 25°F 72°F 84°F 64°F
30% TBP = 10% TBP = IBP TBP = 70% TBP = 90% TBP = EP TBP =
140°F 77°F 26°F 297°F 383°F 501°F
LCO TBP (From Appendix 9: TBP 50% point = 561°F) Given D-86 50% - 30% = 30% - 10% = 10% - IBP = 70%-50% = 90% - 70% = EP-90% =
From Appendix 10 33°F 4FF 73°F 29°F 51°F 39°F
30%TBP = 511°F 10% TBP = 441°F IBP TBP = 343°F 70%TBP = 601°F 90% TBP = 660°F EPTBP = 712°F
Decanted Oil TBP (From Appendix 9: TBP 50% point = 744°F) Given D-86
From Appendix 10
50% 30% 10% 70% 90%
30% TBP = 694°F 10% TBP = 624°F IBP TBP = 425°F 70% TBP = 807°F 90% TBP = 886°F
-
30% = 33°F 10% = 41°F IBP = 236°F 50% = 51°F 70% = 75°F
157
58
Fluid Catalytic Cracking Handbook Table 5-6 Normalized and Adjusted FCC Weight Balance Summary
Stream
Fresh Feed
bpd
ib/hr
Vol% of Feed
Wt% of Feed
50,000
658,814
100.00
100,00
Products
497 7,585 7,549 5,187 20,818
H, C, C,
= Total C2 and lighter
H 2S C,=
C
3
C4 NC4
c;
Gasoline (C5+ to 430°F TBP)
CO (430°F TBP to 650°F TBP)
DO (65Q°F+ TBP)
Coke Total
rue Conversion nerts
0.07 1.15 1.15 0.79 3.16
2,090 4,027 2,064 1,827 4,178
1 ,032 15,262 30,504 16,958 15,579 37,150
4.18 8.05 4.13 3.65 8.36
0.1.6 2 32 4.63 2.57 2.36 5.64
28,597
312,073
57.19
47,37
9,435
1 26,004
18.87
19.13
3,718
55,994
7.44
8.50
55,936
27,440 658,814
111.87
4.17 100.00
73.7
72.3
5,143
HEAT BALANCE
A cat cracker continually adjusts itself to stay in heat balance. This means that the reactor and regenerator heat flows must be equal Figure 5-4). Simply stated, the unit produces and burns enough coke o provide energy to:
Unit Monitoring and Control
Steam
Steam Oil Feed Figure 5-4. Reactor-regenerator heat balance.
159
60
Fluid Catalytic Cracking Handbook
• Increase the temperature of the fresh feed, recycle, and atomizing steam from their preheated states to the reactor temperature « Provide the zendothermic heat of cracking • Increase the temperature of the combustion air from the blower discharge temperature to the regenerator flue gas temperature • Make up for heat losses from the reactor and regenerator to the surroundings • Provide for miscellaneous heat sinks, such as stripping steam and catalyst cooling
A heat balance can be performed around the reactor, around the tripper-regenerator, and as an overall heat balance around the reactoregenerator. The stripper-regenerator heat balance can be used to alculate the catalyst circulation rate and the catalyst-to-oil ratio.
Heat Balance Around Stripper-Regenerator
If a reliable spent catalyst temperature is not available, the stripper s included in the heat balance envelope (II) as shown in Figure 5-4, The combustion of coke in the regenerator satisfies the following eat requirements: « Heat to raise air from the blower discharge temperature to the regenerator dense phase temperature • Heat to desorb the coke from the spent catalyst • Heat to raise the temperature of the stripping steam to the reactor temperature • Heat to raise the coke on the catalyst from the reactor temperature to the regenerator dense phase temperature • Heat to raise the coke products from the regenerator dense temperature to flue gas temperature • Heat to compensate for regenerator heat losses • Heat to raise the spent catalyst from the reactor temperature to the regenerator dense phase temperature
Using the operating data from the case study, Example 5-5 shows eat balance calculations around the stripper-regenerator. The results re used to determine the catalyst circulation rate and the delta coke. Delta coke is the difference between coke on the spent catalyst and oke on the regenerated catalyst.
Unit Monitoring and Control
161
Example 5-5 Stripper-Regenerator Heat Balance Calculations
I. Heat generated in the regenerator: C to CO2 = 24,288 Ib/hr x 14,087 Btu/lb = 342 x 106 Btu/hr H2 to H2O = 2,695 Ib/hr x 51,571 Btu/lb = 139 x 106 Btu/hr S to SO2 = 225 Ib/hr x 3,983 Btu/lb = 0.9 x 106 Btu/hr Total heat released in the regenerator: 342 + 139 + 0.9 = 482 x 106 Btu/hr II. Required heat to increase air temperature from blower discharge to the regenerator dense phase temperature: From Figure 5-5, enthalpies of air at 374°F and at 1,309°F are 90 Btu/lb and 355 Btu/lb. Therefore, the required heat is = 407,493 Ib/hr x (355 - 90) Btu/lb = 108.0 x 106 Btu/hr
III. Energy to desorb coke from the spent catalyst: Desorption of coke = 27,208 Ib/hr x 1,450 Btu/lb = 39.5 x 106 Btu/hr IV. Energy to heat the stripping steam: Enthalpy of 50 psig-saturated steam = 1,179 Btu/lb Enthalpy of 50 psig at 972°F =1,519 Btu/lb Change of enthalpy = 13,000 Ib/hr x (1,519 - 1,179) Btu/lb = 4.4 x 106 Btu/hr V. Energy to heat the coke on the spent catalyst: 27,231 Ibs/hr x 0.4 Btu/lb-°F x (1,309 - 972)°F = 3.7 x 106 Btu/hr
VI. Energy to heat the flue gas from regenerator dense phase to regenerator flue gas temperature: From Figure 5-5, enthalpy of flue gas at 1,309°F = 365 Btu/lb and at 1,330°F = 370 Btu/lb. The required heat is therefore = 433,445 Ib/hr x (370 - 355)°F = 2.6 x 106 Btu/hr
VII, Heat loss to surroundings: Assume heat loss from the stripper-regenerator (due to radiation and convection) is 4% of total heat of combustion, i.e., 0.04 x 482.4 MM Btu/hr = 19.3 x 106 Btu/hr
62
Fluid Catalytic Cracking Handbook
VIII. Energy required to heat the spent catalyst from its reactor to the regenerator temperature = 481.9 - 108.0 - 39.5 - 4.4 - 3.7 - 2.6 - 19.3 = 304.4 x 106 Btti/hr IX. Calculation of catalyst circulation ^ , „. . . Catalyst Circulation =
304.4 xl0 6 Btu/hr (0.285 Btu/°F-lb) x (1,309 - 972)°F
= 3.169 x 106 Ibs/hr = 26.4 short tons/min. Where: 0.285 is the catalyst heat capacity (see Figure 5-6) Cat/oil ratio = 3.169 x 106/658,914 = 4.8 .„ , Coke Yield, wt% 4.2 A 0_ „ ACoke = = — = 0.87 wt% cat/oil ratio 4,8
Reactor Heat Balance
The hot regenerated catalyst supplies the bulk of the heat required o vaporize the liquid feed (and any recycle) to provide the overall ndothermic heat of cracking, and to raise the temperature of disperion steam and inert gases to the reactor temperature. Heat In
Heat Out
Fresh Feed Recycle Air Steam
Reactor Vapors Flue Gas Losses
The calculation of heat balance around the reactor is illustrated in Example 5-6. As shown, the unknown is the heat of reaction. It is alculated as the net heat from the heat balance divided by the feed ow in weight units. This approach to determining the heat of reaction s acceptable for unit monitoring. However, in designing a new cat racker, a correlation is needed to calculate the heat of reaction. The eat of reaction is needed to specify other operating parameters, such
W
I
f
«M#
0J
E
re
to
of the FCC
20
30
40
as a function of the
Alumina Content, Wt.%
50
60
content,
Unit Monitoring and Control
165
as preheat temperature. Depending on conversion level, catalyst type, and feed quality, the heat of reaction can vary from 120 Btu/lb to 220 Btu/lb. In the unit, the heat of reaction is a useful tool. It is an indirect indication of heat balance accuracy. Trending the heat of reaction on a regular basis provides insight into reactions occurring in the riser and the effects of feedstock and catalyst changes. Example 5-6 Reactor Heat Balance
I. Heat into the reactor 1. Heat with regenerator catalyst = 3.169 x 106 Ib/hr x 0.285 Btu/lb-°F x 1,309°F = 1,182.4 x 106 Btu/hr = 1,182.4 x 106 Btu/hr 2. Heat with the fresh feed: At a feed temperature of 594°F, °API gravity = 25.2 and K factor = 12.08, the feed liquid enthalpy is 405 Btu/lb (see Figure 5-7), therefore, heat content of the feed is = 658,914 Ib/hr x 405 Btu/lb = 266.9 x 106 Btu/hr. 3. Heat with atomizing steam: From steam tables, enthalpy of 150 Ib saturated steam = 1,176 Btu/lb, therefore, heat with steam = 10,000 Ib/hr x 1,176 Btu/lb = 11.8 x 106 Btu/hr. 4 Heat of adsorption: The adsorption of coke on the catalyst is an exothermic process; the heat associated with this adsorption is assumed to be the same as desorption of coke in the regenerator (i.e., 35.3 x 106 Btu/hr). Total heat in = 1,182.4 + 266.9 + 11.8 + 35.3 = 1,496.4 x 106 Btu/hr. II. Heat out of the reactor 1, Heat with spent catalyst = 3,169 x 106 Ib/hr x 0.285 Btu/lb-°F x 972°F = 878 x 106 Btu/hr. 2, Heat required to vaporize feed: From Figure 5-8, enthalpy reactor vapors = 778 Btu/lb, therefore, heat content of the vaporized products = 658,814 Ib/hr x 778 Btu/lb = 512.6 x 106 Btu/hr. 3. Heat content of steam: Enthalpy of steam @ 972°F = 1,519 Btu/lb, therefore, heat content of steam = 10,000 Ib/hr x 1,519 Btu/lb = 15.2 x 106 Btu/hr. 4. Heat loss to surroundings: Assume heat loss due to radiant and convection to be 2% of heat with the regenerated catalyst (i.e., 0.02 x 304.4 = 6.1 x 106 Btu/hr)
61
Fluid Catalytic Cracking Handbook
I. Calculation of heat of reaction Total heat out = total heat in Total heat out = 878 x 106 + 512.6 x 106 + 15.2 x 106 + 6.1 x 1C)6 + overall heat of reaction = Total heat in = 1,499.6 x 106Btu/hr Overall endothermic heat of reaction = 84.5 x 106 Btu/hr or —» 128.2 Btu/lb of feed.
Analysis of Results
Once the material and heat balances are complete, a report must be written. It will first present the data. It will then discuss factors ffecting product quality and any abnormal results. It will then discuss he key findings and recommendations to improve unit operation. In the previous examples, the feed characterizing correlations in Chapter 2 are used to determine composition of the feedstock. The esults show that the feedstock is predominantly paraffinic (i.e., 61.6% araffins, 19.9% naphthenes, and 18.5% aromatics). Paraffinic feedocks normally yield the most gasoline with the least octane. This onfirms the relatively high FCC gasoline yield and low octane bserved in the test run. This is the kind of information that should e included in the report. Of course, the effects of other factors, such s catalyst and operating parameters, will also affect the yield structure nd will be discussed. The coke calculation showed the hydrogen content to be 9.9 wt%. As discussed in Chapter 1, every effort should be made to minimize he hydrogen content of the coke entering the regenerator. The hydroen content of a well-stripped catalyst is in the range of 5 wt% to wt%. A 9.9 wt% hydrogen in coke indicates either poor stripper peration and/or erroneous flue gas analysis.
RESSURE BALANCE
Pressure balance deals with the hydraulics of catalyst circulation in he reactor/regenerator circuit. The pressure balance starts with the atic pressures and differential pressures that are measured. The arious pressure increases and decreases in the circuit are then calulated. The object is to:
5 I
920
980
-*-K«11
960
1000
-*-K-t2
Deg. F
1040
-*-K»13
1020
1060
Figyre 5-8. Hydrocarbon vapor enthalpies at various Watson K factors.
940
1080
Unit Monitoring and Control
« * * *
169
Maximize catalyst circulation Ensure steady circulation Maximize the available pressure drop at the slide valves Minimize the loads on the blower and the wet gas compressor
A clear understanding of the pressure balance is extremely important n "squeezing" the most out of a unit. Incremental capacity can come rom increased catalyst circulation or from altering the differential pressure between the reactor-regenerator to "free up" the wet gas compressor or air blower loads. One must know how to manipulate he pressure balance to identify the "true" constraints of the unit. Using the drawing(s) of the reactor-regenerator, the unit engineer must be able to go through the pressure balance and determine whether t makes sense. He or she needs to calculate and estimate pressures, densities, pressure buildup in the standpipes, etc. The potential for mprovements can be substantial.
Basic Fluidization Principals
A fluidized catalyst behaves like a liquid. Catalyst flow occurs in he direction of a lower pressure. The difference in pressure between any two points in a bed is equal to the static head of the bed between hese points, multiplied by the fluidized catalyst density, but only if he catalyst is fluidized. FCC catalyst can be made to flow like a liquid, but only if the pressure force is transmitted through the catalyst particles and not the vessel wall. The catalyst must remain in a fluidized state as it makes a loop through the circuit. To illustrate the application of the above principals, the role of each major component of the circuit is discussed in the following sections, ollowed by an actual case study. As a reference, Appendix 8 contains luidization terms and definitions commonly used in the FCC.
Major Components of the Reactor-Regenerator Circuit
The major components of the reactor-regenerator circuit that either produce or consume pressure are as follows: * Regenerator catalyst hopper * Regenerated catalyst standpipe
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• • • • •
Fluid Catalytic Cracking Handbook
Regenerated catalyst slide (or plug) valve Riser Reactor-stripper Spent catalyst standpipe Spent catalyst slide (or plug) valve
egenerator Catalyst Hopper
In some FCC units, the regenerated catalyst flows through a hopper rior to entering the standpipe. The hopper is usually internal to the egenerator and often of an inverted cone design. It provides sufficient me for the regenerated catalyst to be deaerated before entering the tandpipe. This causes the catalyst entering the standpipe to have maximum flowing density. The higher the density, the greater the ressure buildup in the standpipe. In some FCC designs, the regenerated atalyst hopper is external with fluffing aeration to control the catalyst ensity entering the standpipe.
egenerated Catalyst Standpipe
The standpipe's height provides the driving force for transferring he catalyst from the regenerator to the reactor. The elevation differnce between the standpipe entrance and the slide valve is the source f this pressure buildup. For example, if the height difference is 30 eet (9.2 meters) and the catalyst density is 40 lb/ft3 (641 kg/m3), the ressure buildup is: 40 1h 1 ft2 Pressure Gain = 30 ft x -^3 x , = 8.3 psi (57Fkp) ft 144 in2
The key to obtaining maximum pressure gain is to keep the catalyst uidized over the length of the standpipe. Longer standpipes will equire external aeration. This compensates for compression of the ntrained gas as it travels down the standpipe. Aeration should be dded evenly along the length of the standpipe. In shorter standpipes ufficient flue gas is often carried down with the regenerated catalyst o keep it fluidized and supplemental aeration is unnecessary. Overeration leads to unstable catalyst flow and must be avoided.
Unit Monitoring and Control
171
Aside from proper aeration, the flowing catalyst must contain sufficient 0-40 micron fines to avoid defluidization.
Regenerated Catalyst Slide Valve
The purpose of the regenerated catalyst slide valve is threefold: to regulate the flow of regenerated catalyst to the riser, to maintain pressure head in the standpipe, and to protect the regenerator from a flow reversal. Associated with this control and protection is usually a I psi to 8 psi (7 Kp to 55 Kp) pressure drop across the valve.
Riser
The hot-regenerated catalyst is transported up the riser and into the reactor-stripper. The driving force to carry this mixture of catalyst and vapors comes from a higher pressure at the base of the riser and the ow density of the catalyst/vapor mix. The large density difference between the fluidized catalyst on the regenerator side (approximately 40 b/ft3) and the mixture of cracked hydrocarbon vapors and catalyst on the riser side (approximately 1 lb/ft3) drives the system. As for the pressure balance, this transport of catalyst results in a pressure drop in a range of 5 psi to 9 psi (35 Kp to 62 Kp). This drop is due to static head and, to a esser extent, friction and acceleration of the fluid. In an existing riser, operating changes, such as higher catalyst circulation or lower vapor velocity, can affect the density of reaction mixture and increase pressure drop. This will affect the slide valve differential and percent opening.
Reactor-Stripper The catalyst bed in the reactor-stripper is important for three reasons: « to provide enough residence time for proper stripping of the entrained hydrocarbon vapors prior to entering the regenerator; • to provide adequate static head for flow of the spent catalyst to the regenerator; and • to provide sufficient backpressure to prevent reversal of hot flue gas into the reactor system.
Assuming a stripper with a 20-ft bed level and a catalyst density of 40 lb/ft 3 , the static pressure is: 3 , 40 lbs/ft .. . 20n ft x f— r 2 2 = 5.5 psi
144 in /ft
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Fluid Catalytic Cracking Handbook
pent Catalyst Standpipe
From the bottom of the stripper, the spent catalyst flows into the pent catalyst standpipe. Sometimes the catalyst is partially defluidized n the stripper cone. To counter this, "dry" steam is usually added hrough a distributor) to fluidize the catalyst prior to its entering the tandpipe. The loss of fluidization in the stripper cone can cause a uildup of dense phase catalyst along the cone walls. This buildup can estrict catalyst flow into the standpipe, causing erratic flow and educing pressure buildup in the standpipe. Like the regenerated catalyst standpipe, the spent catalyst standpipe may require supplemental aeration to obtain optimum flow charcteristics. "Dry" stearn is the usual aeration medium.
pent Catalyst Slide or Plug Valve
The spent catalyst slide valve is located at the base of the standpipe. controls the stripper bed level and regulates the flow of spent atalyst into the regenerator. As with the regenerated catalyst slide alve, the catalyst level in the stripper generates pressure as long as is fluidized. The pressure differential across the slide valve will be t the expense of consuming a pressure differential in the range of psi to 6 psi (20 kp to 40 kp). In earlier Model II and Model III FCC units, spent catalyst was ansported into the regenerator using 50% to 100% of combustion ir. This spent cat riser was designed for a minimum air velocity of 0 ft/sec (9.1 m/sec).
Case Study
A survey of the reactor-regenerator circuit of a 50,000 bpd (331 m3/hr) cat cracker produced these results:
eactor dilute phase (dome) pressure eactor catalyst dilute phase bed level eactor-stripper catalyst bed level eactor-stripper catalyst density pent, catalyst standpipe elevation ressure above the spent catalyst slide valve pent catalyst slide valve AP (@ 55% opening)
= = = = = = =
19.0 psig/131 Kp 25.0 ft/7.6 m 18.0 ft/5.5 m 40 Ib/ft3/640 kg/m3 14.4 ft/4.4 m 26.1 psig/180 Kp 4.0 psi/27.6 Kp
Unit Monitoring and Control
Regenerator dilute phase catalyst level Regenerator dense phase catalyst bed level Catalyst density in the regenerator dense phase Regenerated catalyst standpipe elevation Pressure above the regenerated catalyst slide valve Regenerated catalyst slide valve AP (@ 30% opening) Reactor-regenerator pressure AP
173
= 27.0 ft/8.2 m = 15.0 ft/4.6 m = 25 Ib/ft3/400 kg/m3 = 30.0 ft/9.1 m = 30.5 psig/210.3 Kp = 5.5 psi/37.9 Kp = 3 . 0 psi/20.7 Kp
Also, see Figure 5-9 for a graphical representation of the preliminary esults. Starting with the reactor dilute pressure as the working point, the pressure head corresponding to 25 feet (7.6 m) of dilute catalyst ines is: (25 ft) x (0.6 lb/ft3) x (1 ft2/!44 in 2 ) = 0.1 psig (0.7 Kp)
Therefore, the pressure at the top of the stripper bed is: 19.0 + 0.1 = 19.1 psig (131.7 Kp)
The static-pressure head in the stripper is: (18 ft) x (40 lb/ft3) x (1 ft/144 in 2 ) = 5.0 psig (34.5 Kp)
The pressure above the spent catalyst standpipe is: 19.1 + 5.0 = 24.1 psig (166.2 Kp)
The pressure buildup in the spent catalyst standpipe is:
26.1 - 24.1 =2 psi (13.8 K p )
The pressure below the spent catalyst slide valve is: 26.1 -4.0 = 22.1 psig (152 Kp)
The pressure head corresponding to 28 feet (8.5 m) of dilute catalyst ines in the regenerator is: (28 ft) x (1 lb/ft3) x (1 ft2/144 in2) = 0.2 psig (1.4 Kp)
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Fluid Catalytic Cracking Handbook
Rx Vapor
Reactor
Psi diff.
Oil Feed Figure 5-9. Preliminary pressure balance survey.
Unit Monitoring and Control
175
The pressure in the regenerator dome is: 22. J - 0,2 = 21.9 psig (151.0 KP)
The static pressure head in the regenerator is: (18 ft) x (25 lb/ft3) x (1 ft2/!44 in2) = 3.1 psig (21.4 Kp)
The pressure above the regenerated catalyst standpipe is: 22.1 + 3.1 = 25.2 psig (173.7 Kp)
The pressure buildup in the regenerated catalyst standpipe is: 30.5 - 25.2 = 5.3 psi (36.5 Kp)
The pressure below the regenerated catalyst slide valve is: 30.5 - 5.5 = 25 psig (172.4 Kp)
The pressure drop in the vertical riser is: 25 - 19 = 6 psi (41.4 Kp)
The catalyst density in the spent catalyst standpipe is: (2.0 lb/in2) x (144 in2/ft2)/(14.4 ft) = 20 lb/ft3 = 320 kg/m3
The catalyst density in the regenerated catalyst standpipe is: (5.3 lb/in 2 ) x (144 in2/ft2)/(30 ft) = 25.4 lb/ft3 = 407 kg/m3
Figure 5-10 shows the results of the above pressure balance survey.
Analysis of the Findings
The pressure balance survey indicates that neither the spent nor the egenerated catalyst standpipe is generating "optimum" pressure head. This is evidenced by the low catalyst densities of 20 lb/ft3 (320 kg/m3) nd 25.4 lb/ft3 (407 kg/m3), respectively. As indicated in Chapter 8, everal factors can cause low pressure, including "under" or "over"
76
Fluid Catalytic Cracking Handbook
Rx Vapor Reactor
Psi diff.
Oil Feed
igure 5-10. Pressure balance survey with calculated standpipe densities.
Unit Monitoring and Control
177
aeration of the standpipes. In a well-fluidized standpipe, the expected catalyst density is in the range of 35 - 45 lb/ft3 (561 kg/m 3 to 721 kg/m3). If the catalyst density in the spent catalyst standpipe was 40 lb/ft'* (640 kg/m3) instead of 20 lb/ft3 (320 kg/m3), the pressure buildup would have been 4.0 psi instead of 2.0 psi. The extra 2 psi (13.8 Kp) can be used to circulate more catalyst or to lower the reactor pressure. In the regenerated catalyst standpipe, a 40 lb/ft3 (640 kg/m3) catalyst ^ ' 3 density versus a 25.4 lb/ft (407 kg/m") density produces 3 psi (20,7 Kp) more pressure head, again allowing an increase in circulation or a reduction in the regenerator pressure (gaining more combustion air).
Process control instrumentation controls the FCC unit in a safe, monitored mode with limited operator intervention. Two levels of process control are used: • Basic supervisory control • Advanced process control (APC)
Basic Supervisory Control
The primary controls in the reactor-regenerator section are flow, emperature, pressure, and catalyst level. The flow controllers are often used to set desired flows for the fresh eed, stripping steam, and dispersion steam. Each flow controller usually has three modes of control: manual, auto, and cascade. In manual mode, the operator manually opens or closes a valve to the desired percent opening. In auto mode, the operator enters the desired low rate as a set-point. In cascade mode, the controller set-point is an input from another controller. The reactor temperature is controlled by a temperature controller that egulates the regenerated catalyst slide valve. The regenerator temperaure is not automatically controlled but depends on its mode of operation. In partial combustion, the regenerator temperature is conrolled by adjusting the flow of combustion air to the regenerator. In ull burn, the regenerator temperature is a function of operating conditions such as reactor temperature and slurry recycle.
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Fluid Catalytic Cracking Handbook
The reactor pressure is not directly controlled; instead, it floats on he main column overhead receiver. A pressure controller on the verhead receiver controls the wet gas compressor and indirectly ontrols the reactor pressure. The regenerator pressure is often conolled directly by regulating the flue gas slide or butterfly valve. In ome cases, the flue gas slide or butterfly valve is used to control the ifferential pressure between the regenerator and reactor. The reactor or stripper catalyst level controller is controlled with level controller that regulates the movement of the spent catalyst lide valve. The regenerator level is manually controlled to maintain atalyst inventory.
egenerated and Spent Catalyst Slide Valve ow Differential Pressure Override
Normally, the reactor temperature and the stripper level controllers egulate the movement of the regenerated and spent catalyst slide alves. The algorithm of these controllers can drive the valves either ully open or fully closed if the controller set-point is unobtainable. is extremely important that a positive and stable pressure differential e maintained across both the regenerated and spent catalyst slide alves. For safety, a low differential pressure controller overrides the emperature/level controllers should these valves open too much. The hutdown is usually set at 2 psi (14 Kp). The direction of catalyst flow must always be from the regenerator o the reactor and from the reactor back to the regenerator. A negative ifferential pressure across the regenerated catalyst slide valve can llow hydrocarbons to back-flow into the regenerator. This is called flow reversal and can result in an uncontrolled afterburn and ossible equipment damage. A negative pressure differential across he spent catalyst slide valve can allow air to back-flow from the egenerator into the reactor with equally disastrous consequences. To protect the reactor and the regenerator against a flow reversal, ressure differential controllers are used to monitor and control the ifferential pressures across the slide valves. If the differential pressure alls below a minimum set-point, the pressure differential controller PDIC) overrides the process controller and closes the valve. Only fter the PDIC is satisfied will the control of the slide valve return to he process.
Unit Monitoring and Control
179
To maximize the unit's profit, one must operate the unit simulaneously against as many constraints as possible. Examples of these onstraints are limits on the air blower, the wet gas compressor, eactor/regenerator temperatures, slide valve differentials, etc. The onventional regulatory controllers work only one loop at a time and hey do not talk to one another. A skilled operator can "push" the unit gainst more than one constraint at a time, but the constraints change ften. To operate closer to multiple constraints, a number of refiners ave installed an advanced process control (APC) package either within their DCS or in a host computer. The primary advantages of an APC are: * It provides more precise control of the operating variables against the unit's constraints and, therefore, obtains incremental throughput or cracking severity. * It is able to respond quickly to ambient disturbances, such as cold fronts or rainstorms. It can run a day/night operation, taking advantage of the cooler temperatures at night. * It pushes against two or more constraints rather than one single constraint. It can maximize the air blower and wet gas compressor capacities.
As mentioned above, there are two options for installing an APC. One option is to install an APC within the DCS framework, and the ther is to install a multivariable modeling/control package in a host omputer. Each has advantages and disadvantages, as indicated below,
dvantages of Multivariable Modeling and Control
The multivariable modeling/control package is able to hold more ghtly against constraints and recover more quickly from disturbances. his results in an incremental capacity used to justify multivariable ontrol. An extensive test run is necessary to measure the response f unit variables. In APC on DCS framework, the control structure must be designed, onfigured, and programmed for each specific unit. Modifying the ogic can be an agonizing process. Wiring may be necessary. It is ifficult to both document the programming and to test.
80
Fluid Catalytic Cracking Handbook
With a host computer framework, the control package is all in the oftware. Changing the program can still be agonizing, but the program an be tested off-line. There is more flexibility in the computer system, which can be used for many other purposes, including on-line heat nd weight balances.
Disadvantages of Multivariable Modeling and Control
A multivariable model is like a "black box." The constraints go in nd the signals come out. Operators do not trust a system that takes he unit away from them. Successful installations require good training nd continual communication. The operators must know the interconections in the system. The model may need expensive work if changes are made during a urnaround. If the feed gets outside the range the unit was modeled or, results can be at best unpredictable. An upset can happen for which he system was not programmed. The DCS-based APC is installed in a modular form, meaning operators an understand what the controlled variable is tied to more easily. The host computer-based system may have its own problems, includng computer-to-computer data links. In any APC, the operators must be educated and brought into it efore they can use it. The control must be properly designed, meaning he model must be configured and properly "tuned." The operators hould be involved early and all of them should be consulted since ll four shifts may be running the unit differently.
SUMMARY
The only proper method to evaluate the performance of a cat cracker s by conducting a material and heat balance. One balance will tell where the unit is; a series of daily or weekly balances will tell where he unit is going. The heat and weight balance can be used to evaluate revious changes or predict the result of future changes. As discussed n the next chapter, material and heat balances are the foundation for etermining the effects of operating variables. The material balance test run provides a standard and consistent pproach for daily monitoring. It allows for accurate analysis of yields nd trending of unit performance. The reactor effluent can be deter-
Unit Monitoring and Control
181
mined by direct sampling of the reactor overhead line or by conducting unit test run, The heat balance exercise provides a tool for in-depth analysis of he unit operation. Heat balance surveys determine catalyst circulation ate, delta coke, and heat of reaction. The procedures described in this hapter can be easily programmed into a spreadsheet program to alculate the balances on a routine basis. The pressure balance provides an insight into the hydraulics of atalyst circulation. Performing pressure balance surveys will help the nit engineer identify "pinch points." It will also balance two common onstraints: the air blower and the wet gas compressor. Finally, process control systems allow the unit to operate smoothly nd safely. At the next level, an APC package (whether within the DCS ramework or as a host-based multivariable control system) provides more precise control of operating variables against the unit's contraints. It will gain incremental throughput or cracking severity.
REFERENCES
. Davison Div., W.R. Grace & Co., "Cat Cracker Heat and Material Balance Calculations," Grace Davison Catalagmm, No. 59, 1980. . Hsieh, C. R. and English, A. Ar., "Two Sampling Techniques Accurately Evaluate Fluid-Cat-Cracking Products," Oil & Gas Journal, June 23, 1986, pp. 38-43.
CHAPTER 6
Products and Economics
The previous chapters explained the operation of a cat cracker. However, the purpose of the FCC unit is to maximize profitability for he refinery. The cat cracker provides the conversion capacity that very refinery needs to survive. All crudes have heavy gas oils and uel oil; unfortunately, the market for these products has disappeared. FCC economics makes the refinery a viable entity. Over the years, efineries without cat crackers have been shut down because they were ot profitable. Understanding the economics of the unit is as important as undertanding the heat and pressure balance. The dynamics of FCC economics hanges daily and seasonally. It is dependent on market conditions and he availability of feedstocks. The 1990 Clean Air Act Amendment CAAA) has imposed greater restrictions on quality standards for asoline and diesel. The FCC is the major contributor to the gasoline nd diesel pool and is significantly affected by these new regulations. This chapter discusses the factors affecting yields and qualities of FCC product streams. The section on FCC economics describes several ptions that can be used to maximize FCC performance and the efinery's profit margin.
FCC PRODUCTS
The cat cracker converts less valuable gas oils to more valuable roducts. A major objective of most FCC units is to maximize the onversion of gas oil to gasoline and LPG. The products from the cat racker are: • Dry Gas • LPG • Gasoline 182
Products and Economics
• « • •
183
LCO HCO Decanted Oil Coke
Dry Gas
The gas (C2 and lighter) leaving the sponge oil absorber is commonly referred to as dry gas. Its main components are hydrogen, methane, ethane, ethylene, and hydrogen sulfide (H2S). Once the gas s amine-treated for removal of H2S and other acid gases, it is blended nto the refinery fuel gas system. Depending on the volume percent f hydrogen in the dry gas, some refiners recover hydrogen using rocesses such as cryogenics, pressure-swing absorption, or membrane eparation. The recovered hydrogen is often used in hydrotreating. Dry gas is an undesirable by-product of the FCC unit; excessive ields load up the wet gas compressor (WGC) and are often a contraint. The dry gas yield is primarily due to thermal cracking, metals n the feed, and nonselective catalytic cracking. The main factors that ontribute to the increase of dry gas are: • Increase in the concentration of metals (nickel, vanadium, etc.) on the catalyst • Increase in reactor or regenerator temperatures • Increase in the residence time of hydrocarbon vapors in the reactor • Decrease in the performance of the feed nozzles 8 Increase in the aromaticity of the feed
The overhead stream from the debutanizer or stabilizer is a mix of C3's and C4's, usually referred to as LPG (liquefied petroleum gas). t is rich in olefins, propylene, and butylene. These light olefins play n important role in the manufacture of reformulated gasoline (RFC), Depending on the refinery's configuration, the cat cracker's LPG is sed in the following areas: * Chemical sale, where the LPG is separated into C3's and C4's. The C3's are sold as refinery or chemical grade propylene. The C4 olefins are polymerized or alkylated.
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Fluid Catalytic Cracking Handbook
* Direct blending, where the C4's are blended into the refinery's gasoline pool to regulate vapor pressure and to enhance the octane number. However, new gasoline regulations require reduction of the vapor pressure, thus displacing a large volume of C4's for alternative uses. * Alkylation, where the olefins are reacted with isobutane to make a very desirable gasoline blending stock. A Iky late is an attractive blending component because it has no aromatics or sulfur, low vapor pressure, low end point, and high research and motor octane ratings. * MTBE, where isobutylene is reacted with methanol to produce an oxygenate gasoline additive called methyl tertiary butyl ether (MTBE). MTBE is added to gasoline to meet the minimum oxygen requirement for "reformulated" gasoline.
The LPG yield and its olefinicity can be increased by: * Changing to a catalyst, which minimizes "hydrogen transfer" reactions » Increasing the conversion * Decreasing residence time, particularly the amount of time product vapors spend in the reactor housing before entering the main column » Adding ZSM-5 catalyst additive
An FCC catalyst containing zeolite with a low hydrogen transfer ate reduces resaturation of the olefins in the riser. As stated in Chapter , primary cracking products in the riser are highly olefinic. Most of hese olefins are in the gasoline boiling range; the rest appear in the LPG and LCO boiling range. The LPG olefins do not crack further, but they can become saturated y hydrogen transfer. The gasoline and LCO range olefins can be racked again to form gasoline range olefins and LPG olefins. The lefins in the gasoline and LCO range can also cyclize to form ycloparaffins. The cycloparaffins can react through H2 transfer with lefins in the LPG and gasoline to produce aromatics and paraffins. Therefore, a catalyst that inhibits hydrogen transfer reactions will ncrease olefinicity of the LPG, The conversion increase is accomplished by manipulating the folowing operating conditions: * Increasing the reactor temperature. Increasing the reactor temperature beyond the peak gasoline yield results in overcracking
Products and Economics
185
of the gasoline and LCO fractions. The rate of production and olefinicity of the LPG will increase, Increasing feed/catalyst mix zone temperature. Conversion and LPG yield can be increased by injecting a portion of the feed, or naphtha, at an intermediate point in the riser (see Figure 6-1). Splitting or segregation of the feed results in a high-mix zone temperature, producing more LPG and more olefins. This practice
30% of Feed
70% of Feed
Figure 6-1.
A typical feed segregation scheme.
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Fluid Catalytic Cracking Handbook
is particularly useful where the reactor temperature is already maximized due to a metallurgy constraint. » Increasing catalyst to oil ratio. The catalyst to oil ratio can be increased through several knobs including: reducing the FCC feed preheat temperature, optimizing the stripping and dispersion steam rate, and using a catalyst that deposits less coke on the catalyst,
Reduction of the catalyst/hydrocarbon time in the riser, coupled with he elimination of post-riser cracking, reduces the saturation of the already produced olefins" and allows the refiner to increase the eaction severity. The actions enhance the olefin yields and still operate within the wet gas compressor constraints. Elimination of post-riser esidence time (direct connection of the reactor cyclones to the riser) r reducing the temperature in the dilute phase virtually eliminates ndesired thermal and nonselective cracking. This reduces dry gas and iolefin yields. Adding ZSM-5 catalyst additive is another process available to the efiner to boost production of light olefins. ZSM-5 at a typical conentration of 0.5 to 3.0 wt% is used in a number of FCC units to ncrease the gasoline octane and light olefins. As part of the cracking f low octane components in the gasoline, ZSM-5 also makes C3, C4, nd C5 olefins (see Figure 6-2). Paraffinic feedstocks respond the most o ZSM-5 catalyst additive.
Gasoline
FCC gasoline has always been the most valuable product of a cat racker unit. FCC gasoline accounts for about 35 vol% of the total U.S. gasoline pool. Historically, the FCC has been run for maximum asoline yield with the highest octane.
Gasoline Yield For a given feedstock, gasoline yield can be increased by: • Increasing catalyst-to-oil ratio by decreasing the feed preheat temperature • Increasing catalyst activity by increasing fresh catalyst addition or fresh catalyst activity • Increasing gasoline end point by reducing the main column top pumparound rate
Products and Economics
0
5
10
15
ZSM-5 wt% in Catalyst Inventory Figure 6-2. The effect of ZSM-5 on light-ends yield [5].
* Increasing reactor temperature (if the increase does not over-crack the already produced gasoline)
Gasoline Quality
The Clean Air Act Amendment (CAAA), passed in November 1990, as set new quality standards for U.S. gasoline. A complete discussion f the new gasoline formulation requirements can be found in Chaper 10.
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Fluid Catalytic Cracking Handbook
The key components affecting FCC gasoline quality are octane, enzene, and sulfur and are discussed in the following sections.
Octane. An octane number is a quantitative measure of a fuel mixture's resistance to "knocking." The octane number of a particular ample is measured against a standard blend of n-heptane, which has ero octane, and iso-octane, which has 100 octane. The percent of isoctane that produces the same "knock" intensity as the sample is eported as the octane number. Two octane numbers are routinely used to simulate engine perormance: the research octane number (RON) simulates gasoline erformance under low severity (@600 rpm and 120°F (49°C) air emperature), whereas the motor octane number (MON) reflects more evere conditions (@900 rpm and 300°F (149°C) air temperature). At he pump, road octane, which is the average of RON and MON, s reported. Factors affecting gasoline octane are: A. Operating Conditions 1. Reactor Temperature. As a rule, an increase of 18°F (10°C) in the reactor temperature increases the RON by 1.0 and MON by 0.4. However, the MON contribution comes from the aromatic content of the heavy end. Therefore, at high severity, the MON response to the reactor temperature can be greater than 0.4 number per 18°F. 2. Gasoline End Point, The effect of gasoline end point on its octane number depends on the feedstock quality and severity of the operation. At low severity, lowering the end point of a paraffinic feedstock may not impact the octane number; however, reducing gasoline end point produced from a naphthenic or an aromatic feedstock will lower the octane. 3. Gasoline Reid Vapor Pressure (RVP). The RVP of the gasoline is controlled by adding C4's, which increase octane. As a rule, the RON and MON gain 0.3 and 0.2 numbers for a 1.5 psi (10.3 Kp) increase in RVP. B. Feed Quality 1. °API Gravity, The higher the °API gravity, the more paraffins in the feed and the lower the octane (Figure 6-3). 2. K Factor. The higher the K factor, the lower the octane.
Products and Economics
93
92 U
Z
o 91
90 20
22
24
26
Feed Gravity, "API
82
81 Q Z O
s
80
79 20
22
24
26
Feed Gravity, "API Figure 6-3. Feed gravity comparisons (MON and RON) [7].
189
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Fluid Catalytic Cracking Handbook
3, Aniline Point. Feeds with a higher aniline point are less aromatic and more paraffinic. The higher the aniline point, the lower the octane. 4. Sodium. Additive sodium reduces unit conversion and lowers octane (Figure 6-4). C. Catalyst 1. Rare Earth. Increasing the amount of rare earth oxide (REO) on the zeolite decreases the octane (Figure 6-5). 2. Unit Cell Size. Decreasing the unit cell size increases octane (Figure 6-6). 3. Matrix Activity. Increasing the catalyst matrix activity increases the octane. 4. Coke on the Regenerated Catalyst. Increasing the amount of coke on the regenerated catalyst lowers its activity and increases octane.
Benzene. Most of the benzene in the gasoline pool comes from eformate. Reformate, the high-octane blending component from a eformer unit, comprises about 30 vol% of the gasoline pool. Dependng on the reformer feedstock and severity, reformate contains 3 vol% o 5 vol% benzene. FCC gasoline contains 0.5 to 1.3 vol% benzene. Since it accounts or about 35 vol% of the gasoline pool, it is important to know what ffects the cat cracker gasoline benzene levels. The benzene content n the FCC gasoline can be reduced by: • Short contact time in the riser and in the reactor dilute phase • Lower cat-to-oil ratio and lower reactor temperature • A catalyst with less hydrogen transfer
Sulfur. The major source of sulfur in the gasoline pool comes from CC gasoline. Sulfur in FCC gasoline is a strong function of the feed ulfur content (Figure 6-7). Hydrotreating the FCC feedstock reduces ulfur in the feedstock and, consequently, in the gasoline (Figure 6-8). Other factors that can lower sulfur content are: • Lower gasoline end point (see Figure 6-9) • Lower reactor temperature (see Figure 6-10) • Increased matrix activity of the catalyst (text continued on page 195)
Products and Economics RONC vs. SODIUM COMMERICAL DATA
0.40
0.60
EQUILIBRIUM CAT. SODIUM, WT.%
< 80.5 §80.0 -
79.5 79.0 78.5 _
78.0
0.20
0.40
0.80
EQUILIBRIUM CAT SODIUM, WT. %
Figure 6-4.
Effect of sodium on gasoline octane [8J.
191
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Fluid Catalytic Cracking Handbook
84 PILOT PLANT DATA
83
C5-265T/C5-129"C 265-430"F/129-221*C
82
2; 81 o S
80 79
—8
78 77 0 .0
1.0
2.0
3.0
4.0
REO, WT. % Figure 6-5. Effect of fresh REO on MON [9].
82
95
94
81
93
UJ
z 80 O O o: 79 O
91 au on
I 78
89 88
24.20
I
I
I
24.24
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24,20
24.24
24.28
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UNIT CELL SIZE, A
Effects of unit cell size on research and motor octane [10].
Products and Economics
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ke sure instrument dings are correct
Verify Aeration Gas flow to maximize pressure build-up
Improper placement of the Aeration taps
Use Rotameters instead of Restriction Oriffces(RO's)
i
Restriction Orifices are either plugged or improperly sized
Low Catalyst density in the Standpipe
Troubleshooting catalyst circulation.
Check if the Catalyst properties have changed Figured-IB.
Too much, too little or no Aeration Gases
-fluidization of the talyst in the Standpipe
Insufficient pressure build-up in the Standpipe
Low Pressure Upstream of the Slide Valve
r
r
v
\ f
• Adjust the Pumparound rates • Add Top or Side P/A
\^
^/
^.
High delta P across the Main Fractionator
V
J
\(
Refer to 'Coking/Fouling' Troubleshooting Section
\.
J
High Delta P across the Reactor Ovhd. vapor line
Figure8-1 C. Troubleshooting catalyst circulation.
• Add Fins to the Trim Coolers -19 flns per inch * Water Wash the Condensers • Reduce No. of tube passes on the water side * Check pressure drop between Fin-Fans and Trim Coolers
)r
High delta P across the Overhead Condensers
High Pressure Downstream of the Slide Valve
V
\f
V
• Increase Fluffing Gas or Steam to the base of Riser « Replace the Curved section of the Riser
J
High Delta P across the Riser
40
Fluid Catalytic Cracking Handbook
text continued from page 236)
eactor, slide valves are typically operated with a 25% to 60% opening nd a pressure differential of 2 psi to 6 psi (15–40 Kp). Any increase n the throughput or the conversion will either increase the opening, educe the differential, or both.
Causes of Insufficient
Circulation
A lower pressure differential across the slide valve and/or a higher han "normal" slide valve opening are common evidence of a unit's eaching its circulation limitation. The causes of low differential are • Insufficient pressure buildup upstream of the slide valve * Excessive pressure drop downstream of the slide valve
Catalyst circulation can also be limited by mechanical problems with he slide/plug valves. They may have limited travel or will not open ully. This is indicated by the lack of response when an adjustment is made. The differential pressure and the flow will not change. Troublehooting the valves is a function of the valve design and the vendor will supply troubleshooting information,
Low Upstream Pressure Insufficient pressure upstream of the slide valve can be caused by: • Partial plugging in the standpipe * Too little, too much, or no aeration gas either with the catalyst entering the standpipe or along the standpipe
In a properly designed standpipe, the flow of the catalyst develops smooth and uniform static head over the entire length of the standpipe, provided the catalyst entering the standpipe is properly fluidized. This buildup of pressure head provides the necessary driving force for atalyst circulation. However, as the catalyst travels down the standpipe, it can lose fluidity due to compression of interstitial gas being arried with the catalyst. This is particularly true in a long standpipe nd in low-pressure regenerators operating at low pressure. To retain fluidity of the catalyst and to maintain catalyst densities n the 35 to 45 lb/ft3 (560–720 kg/m3) range (the fluid range), many standpipes require external aeration gas to be injected into the down-flowing
Troubleshooting
241
atalyst. The correct quantity of aeration medium and the correct ocation of aeration taps are essential in achieving optimum catalyst ensity (see Figure 8-2). External aeration is not ordinarily needed in hort standpipes (less than 20 feet) (6 m) because sufficient gas is sually drawn into the standpipe to keep the catalyst fully fluidized.
Troubleshooting Upstream Pressure
Effective troubleshooting of the circulation system requires a methdology similar to the procedures outlined in Figure 8–1. Some of the ey steps are as follows: • Obtain a pressure/density profile upstream and downstream of the slide valves. • Verify any changes in catalyst physical properties. • Ensure that the correct amount of aeration gas is injected along the standpipes. One procedure is to vary the aeration flow until the maximum slide valve differential is observed.
Restriction orifices with upstream pressure regulators are frequently mployed to distribute aeration gas into the standpipes. The orifices ..Slip Joint Rotameter or Orifice/Pressure Regulator
Steam or Nitrogen
Figure 8-2.
Typical standpipe aeration.
42
Fluid Catalytic Cracking Handbook
re sized for critical flow so that the constant flow of aeration gas is roportional to upstream pressure regardless of changes in the downtream pressure. Once the unit is running well, it is often assumed that the aeration ystem is sized properly, but changes in the catalyst physical properties nd/or catalyst circulation rate may require a different purge rate. It hould be noted that aeration rate is directly proportional to catalyst irculation rate. Trends of the E-cat properties can indicate changes n the particle size distribution, which may require changes in the eration rate. Restriction orifices could be oversized, undersized, or lugged with catalyst, resulting in over-aeration, under-aeration, or no eration. All these phenomena cause low pressure buildup and low lide valve differential.
High Downstream Pressure
Sometimes insufficient differential across the regenerated catalyst lide valve is not due to inadequate pressure buildup upstream of the alve, but rather due to an increase in pressure downstream of the slide alve. Possible causes of this increased backpressure are an excessive ressure drop in the "Y" or "J-bend" section, riser, reactor cyclones, eactor overhead vapor line, main fractionator, and/or the main fracionator overhead condensing/cooling system. The pressure drop in the "Y" or "J-bend" section could be from mproper fluidization or a flaw in the mechanical design. There are ften fluffing gas distributors in the bottom of the "Y" or along the J-bend" that are designed to promote uniform delivery of the catalyst nto the feed nozzles. Mechanical damage to these distributors or too ttle or too much fluffing gas affect the catalyst density, causing ressure head downstream of the slide valve. The riser pressure drop is related mainly to the catalyst circulation ate and the slip factor. Catalyst circulation rate is largely a function f the oil feed rate, the reactor temperature, and the feed temperature. ncreasing the feed rate, reactor temperature, or lowering the feed emperature will increase the pressure drop across the riser. Slip factor is defined as the ratio of catalyst residence time in the iser to the hydrocarbon vapor residence time. Some of the factors ffecting the slip factor are circulation rate, riser diameter/geometry, nd riser velocity.
Troubleshooting
243
High pressure in the riser could also be due to insufficient fluidizaon gas in the base of the riser. Fluffing gas will vary the catalyst ensity; more fluffing gas lowers the density in the system and the ackpressure on the slide valve. The pressure drop across the reactor cyclones, reactor vapor line, main fractionator, and main column overhead condensing/cooling ystem can be too high. The pressure drop is primarily a function of apor velocity. Any plugging can increase the pressure drop,
roubleshooting Downstream Pressure Use the same procedure as for upstream pressure.
Erratic Circulation
Erratic circulation occurs when the catalyst is not developing a mooth and uniform static head over the entire length of the standpipe. When this happens, the catalyst packs and bridges across the standpipe. ymptoms of erratic circulation include: • • • • • • • •
Severe vibration and movement of the standpipes A noise similar to train chugging Sudden loss of the pressure above the slide valve Fluctuation in the slide valve delta P Ragged reactor temperature and/or stripper level control Pressure swings in the regenerator and the gas plant Cycling of the slide valve Other instrumentation problems
Causes of Erratic Circulation Several factors contribute to erratic circulation. Included are: • A foreign object, such as a piece of refractory, partially obstructing the flow of catalyst in the standpipe. • Improper aeration—either too much or too little. • Changing catalyst properties. This can be due to changes in the fresh catalyst's physical properties and/or malfunctioning of the cyclones.
44
Fluid Catalytic Cracking Handbook
roubleshooting Erratic Circulation To troubleshoot erratic circulation, one must: » Verify that all the instrument readings are "telling the truth" « Verify all the aeration taps are open * Make sure that neither too much nor too little aeration gas is being applied * Verify aeration gas is not wet * Verify that the fresh catalyst properties have not changed * Verify any recent design changes in the standpipes and/or catalyst hopper * Check recycling of the regenerator fines and/or the slurry recycle
Figure 8-3 shows a step-by-step approach to troubleshooting rratic circulation.
CATALYST LOSSES
Catalyst losses will have adverse effects on the unit operation, the nvironment, and operating cost. Catalyst losses appear as excessive
Problem:
Catalyst is not developing a smooth and uniform static head over the entire length of the Standpipe
Ir
Symptom:
Evidence:
Figure 8-3A.
Catalyst packs and bridges across the Standpipe
Severe vibrations and movement of the Standpipes Fluctuation in the Slide Valve delta P A chugging noise similar to "Train" noise Ragged Reactor temperature and/or Stripper level control Sudden loss of pressure above the Slide Valve Pressure swings in the Regenerator and Gas Plant Troubleshooting erratic catalyst circulation.
Troubleshooting
245
ause:
r Improper Aeration Aeration is not adequate to maintain fluidity due to compression
Solution: > Make sure that neither too much nor too little Aeration Gas is being used
Figure 8-3B.
Catalyst is too coarse: Catalyst fines content is too low
Recycling of Regenerator fines and/or Slurry recycle Consider another Catalyst with a different particle size distribution Check Catalyst Iron Content Check properties of recent Catalyst purchase
A foreign object has partially restricted flow of Catalyst in Standpipe
• Use Gamma Ray to confirm plugging • Use high pressure Steam or Nitrogen to dislodge material
Troubleshooting erratic catalyst circulation.
Cause: Catalyst is defluidized to its bulk density; Low Regenerator Pressure
Wet Aeration Air or Steam
Solution: • Increase Regenerator pressure
Figure 8-3C.
Ensure Aeration medium is dry
Recent mechanical revisions to the Unit
Verify the design or\ Hopper and/or Standpipe Plugged vent line on external Hopper Verify amount and orientation of Aeration injection
Troubleshooting erratic catalyst circulation.
4i
Fluid Catalytic Cracking Handbook
arryover to the main fractionator or losses from the regenerator. Evidences of catalyst losses are: * An increase in the ash and BS&W content of the slurry oil * An increase in the recovery of catalyst fines from the electrostatic precipitator or the tertiary separator » An increase in the opacity of the precipitator stack gases * A decrease in the 0 to 40 microns fraction of the equilibrium catalyst or an increase in average particle size * A gradual loss of the catalyst level in the reactor stripper and/or in the regenerator
Causes of Catalyst Losses Common causes of catalyst losses include: * * * *
Changes Changes Changes Changes
in in in in
catalyst properties operating conditions the mechanical condition of the unit operating practice
Changes in the fresh catalyst's physical properties may contribute o catalyst losses. The losses could be due to the fresh catalyst's being soft." "Softness" is evidenced by the quality of the catalyst binder nd the large amount of 0–40 microns. It will increase the attrition endency of the catalyst and thus its losses. Changes in operating parameters also affect catalyst losses. Exmples are: * An increase and/or decrease in catalyst loading to the cyclones * Overloading the cyclones, even at a constant/or higher efficiency, will result in higher catalyst losses * An increase in the feed atomizing and/or stripping steam, causing catalyst attrition and generating fines * An addition of a large amount of steam to the regenerator, particularly to the torch oil nozzles, again causing catalyst attrition
Often, the main cause of catalyst losses is a change in the mechanical ondition of the unit. Examples are: * Trickle valves are either stuck "closed" or "open," possibly due to the hinges being warped or bound. "Warpage" could be due to
Troubleshooting
• • • • • •
• »
247
exposure to high temperatures. Erosion could be due to excessive gas leakage into the diplegs. Trickle valves have fallen off due to inadequate bracing and/or high superficial gas velocity in the regenerator. Holes have formed in the diplegs because of high cyclone velocity or external impingement. Spalled coke or refractory is lodged in the diplegs. This can be caused by improper curing or inadequate refractory supports. Cracks can form in the internal plenum, possibly due to thermal stresses. The dipleg diameter is either too small or too large. Improperly designed, eroded, or even missing restriction orifices used for steam purge or aeration nozzles could cause catalyst attrition. Catalyst attrition is also caused by broken air and stripping steam distributors. Low catalyst level in the regenerator could uncover the diplegs and allow backflow. High catalyst level can prevent the primary cyclones from draining or prevent the trickle valves from operating properly.
Troubleshooting Catalyst Losses
To stop excessive catalyst losses, it should be identified whether the oss is from the reactor or the regenerator. In either case, the following eneral guidelines should be helpful in troubleshooting catalyst losses: « Verify the catalyst bed levels in the stripper and regenerator vessels. • Conduct a single-gauge pressure survey of the reactor-regenerator circuit. Using the results, determine the catalyst density profile. • Plot the physical properties of the equilibrium catalyst. The plotted properties will include particle size distribution and apparent bulk density. The graph confirms any changes in catalyst properties. • Have the lab analyze the "lost" catalyst for particle size distribution. The analysis will provide clues as to the sources and causes of the losses. • Compare the cyclone loading with the design. If the vapor velocity into the reactor cyclones is low, consider adding supplemental steam to the riser. If the mass flow rate is high, consider increasing the feed preheat temperature to reduce catalyst circulation.
48
Fluid Catalytic Cracking Handbook
• Confirm that the restriction orifices used for instrument purges are in proper working condition and that the orifices are not missing. • Consider switching to a harder catalyst. For a short-term solution, if the losses are from the reactor side, consider recycling slurry to the riser. If the catalyst losses are from the regenerator, consider recycling catalyst fines to the unit.
igure 8–4 is a summary of the above discussions.
Nearly every cat cracker experiences some degree of coking/fouling. Coke has been found on the reactor walls, dome, cyclones, overhead apor line, and the slurry bottoms pumparound circuit. Coking and ouling always occur, but they become a problem when they impact hroughput or efficiency.
Evidence of Coking/Fouling Coking/fouling in the reactor and the main column can be detected by: * Cavitation and/or loss of the main column bottoms pumps * Fouling and subsequent loss of heat transfer coefficient in the bottoms pumparound exchange * High pressure drop across the reactor overhead vapor line * Excessive catalyst carryover to the main column
Causes of Coking/Fouling Coke forms in the reactor and main column circuit because of: * * * *
Changes Changes Changes Changes
in in in in
operating parameters catalyst properties feedstock properties mechanical condition of the equipment
Changes in Operating Parameters
The operating conditions of the unit, particularly during startups and eed interruptions, will have a large influence on the formation of coke. oke normally grows wherever there is a cold spot in the reactor ystem. When the temperature of the metal surfaces in the reactor
« O
0
O
5 n
Troubleshooting
249
Problem: Excessive Catalyst loss can cause Unit Shutdown and possible State or Federal Environmental fines
* Increase in Ash and BS&W content of Slurry Bottoms Product * Increase in flue gas Opacity * Loss of Reactor/Regenerator levels * increase in recovery of fines from Electrostatic Precipitator(ESP) * Increase in Regenerator pressure * Decrease in 0–40 microns fractions of E-Cat * Increase in 80+ microns fractions of E-Cat * Change in Catalyst Average Particle Size (APS)
Evidence:
Figure 8-4A.
Changes in Fresh Catalyst properties Fresh Catalyst is "too soft" - has low concentration of I fines
L
• Analyze make-up for PSD & Attrition 1 Plot E-Cat properties - Look for Abnormal Peaks ' Analyze Fines for PSD Consider using a "harder" Catalyst
Troubleshooting catalyst losses.
Changes in Operating conditions j j j i
j
' Increase in Catalyst circulation rate 1 Increase in Cyclone loading 1 Decrease in Cyclone Inlet velocity 1 Increase in Cyclone Gas Outlet velocity 1 Increase in use of Atomizing Steam 1 Increase in Reactor or Regenerator bed levels 1 Water in Steam x ^ Verify Cyclone loading ' Check for any missing RO's Verify accuracy of the Regen. & Rx Catalyst levels: raise or lower Bed level Recycle Slurry to Riser or recovered fines to Regen.
Figure 8-4B.
Changes in Mechanical conditions
• Trickle Valves are either stuck closed or opened 1 Trickle Valves are either warped or eroded ' Trickle Valves have fallen off Holes in the Oiplegs
Cracks in the Plenum 1
Diplegs diameter either "too small" or "too large" Holes in Catalyst Cooler/ Steam Generator "Chunk" of Coke has fallen into Dipleg
Bumping the Regenerator
Troubleshooting catalyst losses.
50
Fluid Catalytic Cracking Handbook
walls and/or the vapor line falls below the dew point of the vapors, ondensation occurs. Condensation and subsequent coke buildup are ue to cooling effects at the surface. A high fractionator bottoms level, a low riser temperature, and a igh residence time in the reactor dome/vapor line are additional perating factors that increase coke buildup. If the main column level ises above the vapor line inlet nozzle, "donut" shaped coke can form t the nozzle entrance. A low reactor temperature may not fully vaporize the feed; unvaorized feed droplets will aggregate to form coke around the feed ozzles on the reactor walls and/or the transfer line. A long residence ime in the reactor and transfer line also accelerate coke buildup. Insufficient bottoms pumparound to the main column heat-transfer one can also form coke.
Changes in Catalyst Properties
Certain catalyst properties appear to increase coke formation. Catalysts with high rare earth content tend to promote hydrogen transfer reacons. Hydrogen transfer reactions are bimolecular reactions that can roduce multi-ring aromatics.
Changes in Feedstock Properties
The quality of the FCC feed also impacts coke buildup in the reactor nternals and vapor line and fouling/coking of the main column circuit. The asphaltene or the resid content of the feed, if not converted in he riser, can contribute to this coking.
Changes in Mechanical Condition of the Equipment
Damaged or partially plugged feed nozzles can contribute to coke ormation due to poor feed atomization. Damaged shed-trays in the bottom section of the main cloumn can ause coke formation due to non-uniform contact between upflowing apors and downflowing liquid.
Troubleshooting Steps
The following are some of the steps that can be taken to minimize oking/fouling:
Troubleshooting
251
« Avoid dead spots. Coke grows wherever there is a cold spot in the system. Use "dry" dome steam to purge hydrocarbons from the stagnant area above the cyclones. Dead spots cause thermal cracking, * Minimize heat losses from the reactor plenum and the transfer line. Heat loss will cause condensation of heavy components of the reaction products. Insulate as much of the system as possible; when insulating flanges, verify that the studs are adequate for the higher temperature. * Improve the feed/catalyst mixing system and maintain a high conversion. A properly designed feed/catalyst injection system, combined with operating at a high conversion, will crack out highboiling feeds that otherwise could be the precursors for the formation of coke. * Follow proper start-up procedures. Introduce feed to the riser only when the reactor system is adequately heated up. Local cold spots cause coke to build up in the reactor cyclones, the plenum chamber, or the vapor line. * Keep the tube velocity in the bottoms pumparound exchanger(s) greater than 7 ft/sec. Putting the bundles in parallel for more heat recovery may lead to low velocity. * Hold the main column shed tray's liquid temperature under 700°F and minimize the level and residence time of the hot liquid. Ensure adequate wash to shed decks to minimize coking in the bottom of the main column. Some paraffinic feeds may require a lower temperature. * Utilize a continuous-cycle oil flush into the inlet of the bottoms exchanger. This keeps the asphaltenes in solution and increases tube velocity. * Verify that no fresh feed is entering the main column. Feed can enter the main column through emergency bypasses or through the feed surge tank vent line.
igure 8-5 is a summary of the above discussions.
LOW REVERSAL
A stable pressure differential must be maintained across the slide alves. The direction of catalyst flow must always be from the regenerator (text continued on page 254)
Figure 8-5A.
r temperature ence Time in the d Main Column ms temperature s Pumparound Rate nger tube wall e J
in the Reactor heat-up olumn Bottoms
Troubleshooting coking/fouling.
Cracked Feedstock
High Mole weight Asphaltene & Resins, precipitate and bind to process equip. High Levels of
Damaged or pa plugged Feed N Loss of the She Feed leaking th Bottoms exchan Feed Diversion
Changes in Mecha Conditions of the Equipment
Higher pressure drop across the Reactor Overhead vapor line
Changes in feedstock Properties
Fouling and loss of Heat Transfer in Bottoms Exchanger
High Level of RareEarth in the Catalyst Low Catalyst Micro Activity Test (MAT)
Cavitation and/or loss of Main Column Bottoms Pumps
in Operating s
:
Unscheduled Unit interruptions loss of profit and higher maintenance costs
* Increase Bottoms traffic * Inject a continuous Cycle Oil flush into inlet Bottoms PA Exchangers * Install duplex filters upstream of Bottoms Pumps * Install high efficiency feed nozzles * Use 1" or larger tube diameter * Keep C7 insolubles in Slurry System less than 5% wt * Use U-tube for Bottoms Exch. * Draw more Bottoms Product * Have a spare Bottoms Exchanger bundle
Figure 8-5B. Troubleshooting coking/fouling.
comes to Coking, Residence Time is e as increasing temp, by 25°F
ber:
erly insulate RX Overhead piping Main Column Inlet nozzle the tube velocity > 7 ft/sec Main Column Bottoms erature < 700°F a "dry" Dome Steam System
Recommendations:
54
Fluid Catalytic Cracking Handbook
ext continued from page 251)
o the reactor and from the reactor-stripper back to the regenerator. A egative differential pressure across the regenerated catalyst slide valve an allow fresh feed and oil-soaked catalyst to backflow from the riser nto the regenerator. This flow reversal can result in uncontrolled urning in the regenerator and potentially damage regenerator internals, osting a refiner several million dollars in production loss and mainenance expense. Similarly, a negative pressure differential across the spent catalyst lide valve can allow hot flue gas to backflow to the reactor and he main fractionator, severely damaging the mechanical integrity of hese vessels. Some of the main causes of loss of pressure differential across the lide valves are as follows: • Loss of the air blower or the wet gas compressor • Presence of water in the feed • High catalyst circulation rates resulting in excessive slide valve opening and low differential • Loss of regenerator or stripper bed levels • Failure of the reactor temperature controller and reactor-stripper level controller • Bypass open around a shutdown valve Troubleshooting flow reversal is outlined in Figure 8-6.
Reversal Prevention Philosophy
The FCC process is very complex and many scenarios can upset perations. If the upset condition is not corrected or controlled, each cenario could lead to a reversal. Table 8-1 contains a cause/effect hutdown matrix indicating scenarios in which a shutdown (reversal) ould take place. In most cases, a unit shutdown is not necessary if dequate warning (low alarms before low/low shutdowns) is provided. he operating staff must be trained to respond to these warnings. The shutdown system will have adequate interlocks to prevent nadvertent trips. The system must include two-out-of-three voting or ackup instruments. The operators must trust the system for it to emain in service.
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Changes in Catalyst Conditions
I.ower Gasoline Yield
Fig. 8-11 B
[
f
»
\t
•Check pressure profil around Feed Nozzles •Track H2 in Coke •Survey ttie Stripper >-
f
V
feed nozzles •Damaged Stripper Steam Distributor
\f
c •Damaged or plugged
Changes in Mechanical Conditio
^
Lower Gasoline Octa Figure 8-11D
Reduce Gasoline Olefins Figure 8-11E
High Dry Gas Yield
Loss of Revenue Off-Spec Products
Figure 8-11 A. Troubleshooting desired product quantity and quality.
w Fee^ erties k for changes in gen, Nl, V, °F API gravity, K or, Rl & 650°F
e Contaminants s Paraffins e Aromatics of run for Feed
* e Residue
•
Changes in ed Properties
^
er Conversion
Problem:
Troubleshooting
267
The decreases in microactivity and surface area are strong functions f thermal deactivation in the regenerator and the presence of metals n the feed.
Operating Variables The following operating parameters lower conversion: • • • •
Decrease Decrease Decrease Decrease
in in in in
the the the the
reactor temperature catalyst-to-oil ratio atomizing steam fresh catalyst addition rate
Mechanical Conditions
Damaged or plugged feed nozzle(s) and/or damaged stripping steam istributor(s) are the common causes of mechanical failures that ffect "true" conversion. Note that the "apparent" conversion, as iscussed in Chapter 5, is affected by the distillation cut point and main column operations.
roubleshooting Steps • Trend the feedstock properties; look for changes in the K factor, 1,050°F+ (565°C+), aniline point, refractive index, and °API gravity. The feed endpoint may have been increased to fill the unit. The conversion penalty may be a small price to pay for the increased capacity, but the penalty can be minimized. Verify that the refinery LP reflects current data on yields and product quality. • Plot properties of the fresh and equilibrium catalysts; ensure that the catalyst vendor is meeting the agreed quality control specifications. Verify that the catalyst vendor has the latest data on feed properties, unit condition, and target products. Verify the fresh makeup rate. Check for recent temperature excursions in the regenerator or afterburning problems. • Trend the reactor temperature, cat-to-oil ratio, and atomizing steam rate. Verify the accuracy of the reactor temperature thermocouple and atomizing steam flow meter. • Perform a single-gauge pressure survey around the feed nozzles. Calculate the hydrogen content of the spent catalyst. Conduct a
268
Fluid Catalytic Cracking Handbook
gamma ray scan test to verify the mechanical condition of the stripping steam distributor.
Observing a High Dry Gas Yield
Dry gas yield is affected by everything that affects conversion Figure 8-11B). Changes to increase conversion can increase the dry gas yield. High gas yield shows up as higher speed on the compressor (if centrifugal). In many cases, lower molecular weight (due to higher hydrogen content) can have the same effect.
Feedstock Quality The feed parameters that increase the dry gas yield are: * Increase in nickel and vanadium content * Increase in naphthene, olefin, and aromatic concentration, which is indicated by an increase in the refractive index and decreases in aniline point and K factor
Catalyst Properties The E-cat properties that increase dry gas yield are: * Increase in the level of nickel, vanadium, and sodium * Decrease in E-cat activity, surface area, fresh catalyst activity, and rare earth content * Increase in the gas and coke factors of the E-cat
Operating Variables Operating parameters that increase dry gas yield are: * Increase in the reactor temperature * Increase in the regenerator temperature « Decrease in the atomizing steam * Increase in slurry or HCO recycle
Mechanical Conditions Mechanical conditions that can increase dry gas yield are: » A failing reactor temperature thermocouple * Partially plugged or damaged feed nozzles
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CH 3 — CH — CH — CH 2 —CH 3 CH3
PROPYLENE
ypical Yield:
+
ISOBUTANE
CH3
CH3
-> DIMETHYLPENTANE
22
Fluid Catalytic Cracking Handbook
.0 Volume of propylene + 1.3 volume of isobutane -> 1.80 volume f alky late.
Butyiene Alk\ lation CH3 CH3 — C = CH2 + CH3 — CH — CH2 -> CH3 — CH — CH, — CH — CH3 CH,
PROPYLENE +
CH3
CH,
ISOBUTANE
-»
CH3
2,2,4 TRIMETHYLPENTANE
ypical Yield:
.0 volume of butylene +1.2 volume of isobutane —> 1.70 volume f alkylate.
Example 10-2 Etherification of Isobutviene
CH3 — C — CH3 + CH3 — OH CH?
-> CH3 — C — O — CH3 CH3
ISOBUTYLENE + METHANOL -> METHYL TERTIARY BUTYL ETHER (MTBE)
ypical Yield:
.0 volume of isobutane + 0.43 volume of methanol —> 1.27 volume f MTBE.
There are etherification processes, such as MTBE and TAME, aimed t producing ethers from C5, C6, and C7 tertiary olefins. Both alkylate and ether have excellent properties as gasoline blending omponents. They have a low RVP, a high road octane, no aromatics, nd virtually zero sulfur. The emphasis on alkylation and etherification will continue in both the U.S. and the rest of the world.
Emerging Trends in Fluidized Catalytic Cracking
323
A conventional FCC unit can be an "olefin machine" with proper perating conditions and hardware. Catalysts with a low unit cell size and high silica/alumina ratio favor olefins. Additionally, the addition of ZSM, with its lower acid site density and very high framework silica-alumina atio, converts C?+ gasoline into olefins. A high reactor temperature nd elimination of the post-riser residence time will also produce more lefins. Mechanical modification of the FCC riser for "millisecond" racking has shown potential for maximizing olefin yield.
Challenges Facing RFG
RFG is a cost-effective fuel that improves air quality and is a mechanism through which the refining industry can be competitive. The Complex Model is most likely here to stay. The concentration of asoline sulfur must be reduced and the gasoline RVP will most likely e limited to about 7 psi. Nevertheless, in the years to come, numerous ssues regarding RFG will be facing refiners. Most are regulatory, olitical, and bureaucratic issues. Following are some of these issues: « Public perception of RFG regarding health effects of ethers, price increase, and engine performance complaints * EPA's ethanol mandate and the subsequent stay of that mandate by federal court * Complexity of testing, distribution, storage, handling, and blending facilities * Record-keeping and development of a uniform certification program, * Intel-changeability of MTBE to ETBE * Interpreting the baseline * The future of opt-in areas: the continual decline in air quality where RFG is not sold * Antidumping, credits, and trading * The program length of oxygenated fuels for CO nonattainment areas * The definition of "domestic supply"
RESIDUAL FLUIDIZED CATALYTIC CRACKING (RFCC)
Deterioration in the worldwide crude oil supply (Table 10-6), ontinual decline in the demand for heavy fuel oil, and recent mechancal and catalyst advances have provided economic incentives to
24
Fluid Catalytic Cracking Handbook Table 10-6 U.S. Crude Characteristics
Year
°APi Gravity
Wt% Sulfur
1983 1984 1985 1986 1987 1988 1989 1990 1991 1992 1993
32.92 32.96 32.46 32.33 32.22 31.93 32.14 31.86 31.64 31.32 31.30
0.88 0.94 0.91 0.96 0.99 1.04 1.06 1. 10 1.13 1.16 1.15
Source: Swain [24]
upgrade the atmospheric and/or vacuum bottoms in the residual fluidzed catalytic cracking (RFCC) unit. Although residue upgrading in he United States is mostly delayed coker based, most new FCC units are either residue crackers or have in-place provisions to process esidue at a later date. This is more pronounced in the new units built n the Far East, Europe, and Australia. The residue from their crude oils is more paraffinic and contains less metals than North Sea or Middle Eastern crude oils, which makes them more suitable for RFCC. An RFCC is distinguished from a conventional vacuum gas oil FCC n the quality of the feedstock. The residue feed has a high coking endency and an elevated concentration of contaminants.
Coking Tendency
Residue feedstocks have a higher coking tendency, which is indicated by higher levels of Conradson carbon and a higher boiling point. The common definition of residue is the fraction of the feed that boils above 1,050°F and Conradson carbon levels greater than 0.5 wt%. The esidual portion of the feed contains hydrogen-deficient asphaltenes and polynuclear compounds. Some of these compounds will lay down on active catalyst sites as coke, reducing catalyst activity and selectivity.
Emerging Trends in Fluidized Catalytic Cracking
325
Feed Contaminants
The residual portion of feedstocks contains a large concentration of ontaminants. The major contaminants, mostly organic in nature, nclude nickel, vanadium, nitrogen, and sulfur. Nickel, vanadium, and odium are deposited quantitatively on the catalyst. This deposition oisons the catalyst permanently, accelerating production of coke and ight gases. Nickel in the feed is deposited on the surface of the catalyst, romoting undesirable dehydrogenation and condensation reactions. These nonselective reactions increase gas and coke production at he expense of gasoline and other valuable liquid products. The eleterious effects of nickel poisoning can be reduced by the use of ntimony passivation. Vanadium in the feed poisons the FCC catalyst when it is deposited n the catalyst as coke by vanadyl porphydrine in the feed. During egeneration, this coke is burned off and vanadium is oxidized to a V*5 oxidation state. The vanadium oxide (V2O5) reacts with water vapor in the regenerator to vanadic acid, H3VO4. Vanadic acid is mobile and it destroys zeolite crystal through acid-catalyzed hydrolysis. Vanadic acid formation is related to the steam and oxygen conentration in the regenerator. Vanadium and sodium neutralize catalyst acid sites and can cause ollapse of the zeolite structure. Figure 10-5 shows the deactivation f the catalyst activity as a function of vanadium concentration. Destruction of the zeolite by vanadium takes place in the regenerator where the combination of oxygen, steam, and high temperature forms anadic acid according to the following equations: 4V + 5 O2 --» 2 V2O5 V2CL + 3 H2O -* 2 VO (OH)3
The produced vanadic acid, VO (OH)3, is mobile. Sodium tends to ccelerate the migration of vanadium into the zeolite. This acid attacks he catalyst, causing collapse of the zeolite pore structure. The presence of increased basic nitrogen compounds, such as pyridines nd quinoline in the FCC feedstock, also attack catalyst acid sites. The esult is a temporary loss of catalyst activity and a subsequent increase
26
Fluid Catalytic Cracking Handbook
68
^>T~^ - ^ T
67
* *
i ^***^***«*^
"
I
C'C^L
^'rC? N „
>7, ^^^^S-^4
? 65 > 1 64
"j-. j Q $^ ^"""P"-^
a
o 63 5 6?
^T*—^^
I
'~~\?°-o . I * ^ sit^O-.
i "''"i. f
h
t
61
~' ~ '"-*.' '"ft "
~~1
60
0
1000
2000
3000
4000
5000
6000
Vanadium, ppm
Figure 10-5. Vanadium deactivation varies with regenerator severity {25],
n coke and gas yields. Additionally, in the regenerator, some of the adsorbed nitrogen is converted to nitrogen oxide (NOX). Although an increase in the sulfur content of the residue feedstock will have a minimal effect on unit yields, the sulfur content of the RFCC products and the flue gas is greater, requiring additional treatng facilities.
Operational Impacts of Residue Feedstocks In the unit, residue feedstocks have the following effects: * Higher delta coke and coke yield, which are associated with residue feedstocks, will result in elevated regenerator temperature and higher combustion air requirements. * Exposure of the catalyst to a variety of feed contaminants and the higher regenerator temperature will reduce both selectivity and activity. * Greater levels of nitrogen and sulfur in the residue feed increase emissions of NOX and SOX from the regenerator.
Emerging Trends in Fluidized Catalytic Cracking
327
Minimizing Detrimental Effects of Processing Residual Feeds
The proper choice of a feed injection system, regenerator, and atalyst are some of the key aspects of successful RFCC operation. An efficient feed injection system produces extremely small droplets hat vaporize quickly. Rapid vaporization minimizes the amount of on-vaporized hydrocarbons that block the active sites. An effective eed nozzle system must instantaneously vaporize and crack asphaltenes nd poly nuclear aromatics to lower boiling entities. The regenerator design, either single-stage or two-stage, should rovide uniform catalyst regeneration, increase flexibility for processing variety of feedstocks, and minimize thermal and hydrothermal eactivation of the catalyst. The catalyst design should be optimized to achieve the followng objectives: » • • • •
Low coke and gas production Efficient bottoms cracking Improved metals resistance Improved thermal and hydrothermal stability An active matrix and a low hydrogen transfer activity to convert the bottoms and minimize delta coke
REDUCING FCC EMISSIONS
The gaseous emissions from the FCC unit are CO, NOX, pariculates, and SOX, All are either locally or nationally regulated. Table 10-7 shows the current allowable limits of the EPA New Source Performance Standards (NSPS) for the emissions of these irborne pollutants. NSPS levels can be triggered by one of the ollowing conditions: • Construction of a new unit • Revamp of the regenerator, provided the modification costs are more than 50% of a comparable regenerator 8 Any capital modification of the unit that increases its emission rates
There is no national requirement limiting NOX emissions from the CC flue gas, but several state and regional agencies have imposed mits on their release. These emissions are directly proportional to
28
Fluid Catalytic Cracking Handbook Table 10-7 EPA's New Source Performance Standards (NSPS) for Gaseous Emissions from the FCC Regenerators
Source
Allowable Limits
Carbon monoxide (CO)*
Less than 500 ppmv in the flue gas
Nitrogen oxides (NOX)
None (local and regional only)
Participates**
A maximum of 1.0 pound of solids in the flue gas per 1,000 pounds of coke burned
Sulfur oxides (SO2 + SO3)*
Exempt if the feed sulfur is less than 0.30 wt% If there is no add-on control such as a wet gas scrubber, 9.8 kilograms of (SO2 + SO3) per 1,000 kilograms of coke burned. This is approximately equal to 500 ppmv. Add-on device: reduce (SO2 + SO3) by at least 90% or no more than 500 ppmv, whichever is less stringent.
Effective January 1984 *Effective June 1973
he quality of FCC stocks, operating conditions, catalyst type, and mechanical condition of the unit. Processing feeds that contain a high oncentration of residue, sulfur, nitrogen, and metals will release a greater amount of SOX, NOX, and particulates. Various technologies re available to reduce flue gas emissions.
Particulates
Electrostatic precipitators (ESP) and wet gas scrubbers (WGS) are widely used to remove particulates from the FCC flue gas. Both can ecover over 80% of filtrable solids. An ESP (Figure 10-6) is typically nstalled downstream of the flue gas heat recovery (prior to atmosheric discharge) to minimize particulate concentration. If both low articulate and low SOX requirements are to be met, a wet gas scrubber uch as Belco's (Figure 10-7) should be considered. If SOX removal
ORMER IER
EN
ROOF
COLLECTING SURFACE RAPPER
HIGH-VOLTAGE SYSTEM SUPPORT INSULATOR
RAPPER INSULATOR
DISCHARGE ELECTRODE RAPPER
Figure 10-6. Typical electrostatic precipitator (ESP),
HOPPER-
DISCHARGE ELECTRODE
INSULATOR COMPARTMENT,
BUS DUCT-
DO
•SIDE
Figure 10-7. Schematic of Belco scrubbing system (courtesy of Belco Corporation).
CULATING PUMP
SODA
Emerging Trends in Fluidized Catalytic Cracking
331
s not a prime objective, an ESP will be less expensive from the tandpoints of both initial capital and operating costs. In some cases, bag house system can be used instead of an ESP,
OX
Three methods are widely used to reduce SOX emissions from the CC flue gas: FCC feed pretreatment Catalyst additives Flue gas desulfurization
Feed hydrotreating or hydrocracking reduces SOX emissions and the ulfur content of FCC products. As discussed earlier in this chapter, many benefits are associated with FCC feed hydrotreating. It is mportant to note that most of the sulfur in a hydrotreated feed is in eavy organic compounds and will be concentrated in the decanted il and coke. Consequently, for a given sulfur in the feed, more SOX will be produced with hydrotreated feed. For refiners having low to moderate levels of SOX in their FCC flue as (less than 1,000 ppm), SOX additives are usually the most ecoomical method of reducing SOX emissions. These additives are njected separately into the regenerator. They capture SO3 in the egenerator (oxidizing atmosphere) and release sulfur as H2S in the reactor educing atmosphere). A reliable on-line SO2 analyzer will ensure that a ufficient quantity of additive is injected. Operating conditions of the egenerator, especially partial versus full combustion and excess oxygen evel, will greatly influence the additive's effectiveness. When processing high-sulfur feeds (greater than 1.0 wt%) or if the equired SOX reduction levels are greater than 80%, other capitalntensive desulfurization technologies must be considered. Several ue gas desulfurization technologies are available. Haldor Topsoe's WSA, United Engineers' Mgo., Exxon, and Belco (Figure 10-7) wet as scrubbing (WGS) are among the most widely used processes to emove SOX. The WGS process removes both SOX and particulates.
CO
The CO levels released from the regenerator flue gas operating ither in complete or partial combustion are normally less than 10
32
Fluid Catalytic Cracking Handbook
pm. For units operating in partial combustion, the flue gas must e sent to a CO boiler. For units operating in complete combustion, he concentration of CO largely depends on the operating conditions f the regenerator (mainly temperature and excess oxygen), the CO romoter level, and the efficiency of the air/spent catalyst distriution system.
NOX
NOX levels in the FCC flue gas typically range from 50-500 ppm. Nitrogen content of the feed, excess oxygen, regenerator residence ime, dense phase temperature, and CO promoter all influence the oncentration of NOX. In the regenerator, most of the NOX is formed as NO, with little N2O or NO2. About 90% of organic nitrogen in the spent catalyst is onverted to inorganic nitrogen, and a very small amount becomes NO. NO can be lowered by reducing excess oxygen and CO promoter. The resent platinum-based promoter oxidizes intermediates such as HCN nd NH3 to NO and decreases the reducing agent such as CO. To reduce nitrogen oxide, thermal and catalytic processes are availble. The thermal process is licensed by Exxon. NH3 or urea is injected nto the flue gas at an elevated temperature (-1600°F, 870°C); NOX s reduced to nitrogen. This process is applicable to FCC units that ave CO boilers. NOX can also be reduced over a catalyst at 500°F o 750°F (260°C to 400°C).
EMERGING DEVELPMENTS IN CATALYSTS, PROCESSES, AND HARDWARE
The FCC process has a long history of innovation and will continue o play a key role in the overall success of the refining industry. The ontinuing developments will primarily be in the areas of catalyst, rocess, and hardware technologies.
Catalyst
Since the mid-1960s, formulation of FCC catalysts has improved teadily. The focus of the research is in the following areas:
Emerging Trends in Fluidized Catalytic Cracking
* * • • * •
333
Improvement in zeolite quality Improvement in the catalyst's binder properties Increase in the quantity and choice of active matrix Customization of catalyst to the unit's objectives and constraints A widespread use of ZSM-5 or similar zeolite Improvements in the developments of catalyst additives for reducing gasoline sulfur and NOX emission
here has also been an ongoing trend to formulate a higher-quality eolite. Higher quality has been reflected in: • Greater silica-to-alumina (SAR) of zeolite. Greater SAR results in a zeolite that is more stable, yields more olefins, improves octane, and increases product selectivity. * Improved crystallinity by producing more uniform zeolite crystals, FCC catalyst manufacturers have greater control over the zeolite acid site distribution. In addition, there is an upward trend in the quantity of zeolite being included in the catalyst.
The selectivity and activity of the catalyst matrix will continue to mprove. The emphasis on bottoms cracking and steady reduction in he reaction residence time demands an increase in the quantity of ctive matrix. The improvements in the catalyst's binder properties will reduce he catalyst attrition rate; thus, lowering the flue gas stack opacity. his improvement allows refiners to use a "harder" catalyst without dversely affecting the catalyst's fluidization properties. Future catalyst formulation will be customized to meet the individual efiner's needs. Catalyst manufacturers will be tailoring catalysts to meet each refiner's requirements. The demand for ZSM-5 additives will increase because of their nherent ability to crack low-octane, straight chain olefins to C3 and C4 olefins and also to isomerize low-octane linear olefins to higher ctane branched olefins. Once ZSM-5's patent has expired, its use hould increase. Further developments in the effectiveness of the FCC gasoline sulfur eduction additives will allow a number of refiners to meet the equired reduction in gasoline sulfur without undertaking costly capital rojects. Additionally, improvements in the CO promoter additives will reduce NOX emissions when the promoter is used. Finally, other
34
Fluid Catalytic Cracking Handbook
ost-effective additives will be developed to not only reduce NOX missions, but also reduce catalyst related fouling in the regenerator ue gas heat recovery system.
Operating Conditions
FCC will still play a dominant role in producing cleaner-burning uels. The inherent flexibility of the process will allow refiners to meet he fuel reformulation requirements. With the anticipated growing emand for alky late and ethers, the FCC operating parameters will e adjusted to maximize production of propylene, isobutylene, and soamylene. The projected trend in operating conditions will be to a igher reactor temperature, a higher catalyst-to-oil ratio, a higher eaction mix temperature, and shorter catalyst contact time.
Technology Development
Since 1942, when the first FCC unit came onstream, new techologies have continuously evolved to maximize performance to meet he ever-changing product requirements and feedstock qualities. Future echnology development will remain dynamic. Examples of the new nd ongoing technologies aimed at enhancing the unit's operational nd mechanical performance, as well as complying with environmental egulations, are: • Reducing sulfur and aromatics in gasoline and distillate. • Minimizing disposal of equilibrium catalyst. • Minimizing catalyst back-mixing in the riser to minimize production of undesirable products. Redesign of the conventional riser for a down-flow of catalyst and vapors could virtually eliminate back-mixing. • Achieving an ultra-short catalyst-hydrocarbon contact time, designed to maximize olefins and gasoline yields while minimizing the bottoms yields. • Eliminating long dilute-phase residence time downstream of the riser to prevent recracking of hydrocarbon vapors in the reactor housing. • Improving feed and catalyst injection systems. • Improving spent catalyst distribution.
Emerging Trends in Fluidized Catalytic Cracking
335
• Improving mechanical reliability of the FCC reactor-regenerator components. • Increasing use of feed segregation to maximize production of light olefins. « Increasing use of riser quench to maximize the reaction mix temperature and to promote maximum vaporization of the feedstock. • Increasing use of catalyst additives to reduce gaseous emissions and to maximize light olefins. These are just some of the many challenges facing FCC operaons today.
SUMMARY
The United States refining industry is undergoing a restructuring hase. Refiners will continue to be under pressure and only the most fficient and profitable operations are going to survive. The survivors will be those who have some niche in the market place, have the ersatility to handle low-cost crude, meet product demand, and conorm to environmental regulations. FCC is one of the cheapest conversion processes. Its inherent exibility can assist a refiner in meeting changing product requirements in spite of the steady decline in feedstock quality. The U.S. Federal RFG program has imposed new challenges for the CC, particularly regarding the sulfur, aromatics, and olefin content of asoline. Various commercially proven technologies, along with evolving echnologies, will be available to comply with these new rules. The use of RFCC will continue to grow, particularly in regions of he world where atmospheric or vacuum residue contains low levels f contaminants. Careful regenerator and feed injection designs are mportant in ensuring a successful operation. Gaseous emissions (CO, NOX, SOX, particulates) have been regulated t local and national levels. The quantity of these emissions is directly elated to the quality of the FCC stocks, operating conditions, catalyst ype, and mechanical conditions of the unit. Processing heavy feeds will release a greater amount of SOX, NOX, and particulates. In conclusion, FCC has had a long history of innovations. New echnological developments will continue to emerge, optimizing its erformance. Its versatility and high degree of efficiency will continue o play a key role in meeting future market demands.
3§
Fluid Catalytic Cracking Handbook
REFERENCES
1. Mauleon, J. L. and Letzsch, W. S., "The Influence of Catalyst on the Resid FCCU Heat Balance," presented at Katalistik's 5th Annual FCC Symposium, Vienna, Austria, May 23-24, 1984. 2. A. W. Peters, G. Yaluris, G. D. Weatherbee, X. Zhao, "Origin and Control of NOX in the FCCU Regenerator," Grace Davison, Columbia, MD. 3. Davis, K., and Ritter, R. E., "FCC Catalyst Design Considerations for Resid Processing—Part 2," Grace Davison Catalagram, No. 78, 1988. 4. Hammershaimb, H. U., and Lomas, D. A., "Application of FCC Technology to Today's Refineries," presented at Katalistiks' 6th Annual FCC Symposium, Munich, Germany, May 22-23, 1985. 5. Kool, J. M., "Commercial Experience with Resid Cracking in Conventional FCC Units," presented at the 1984 Akzo Chemicals Symposium, 6. Hood, R., and Bonilla, J., "Residue Upgrading by Solvent Deasphalting and FCC," presented at the Stone & Webster 5th Annual Meeting, Dallas, Texas, October 12, 1993. 7. Dean, R. R., Kibble, P. W., and Brown, G. W., "Crude Oil Upgrading Utilizing Residual Oil Fluid Catalyst Cracking," presented at Katalistiks' 8th Annual FCC Symposium, Budapest, Hungary, June 1-4, 1987. 8. Johnson, T. E., "Resid FCC Regenerator Design," presented at the M.W. Kellogg Co. Refiing Technology seminar, Houston, Texas, February 9-10, 1995. 9. Letzsch, W., Mauleon, J. L. Jones, G., and Dean, R., "Advanced Residual Fluid Catalytic Cracking," presented at Katalistiks' 4th Annual FCC Symposium, Amsterdam, The Netherlands, May 18-19, 1983. 10. Elvin, F. J., and Krikorian, K. V., "The Key to Residue Cracking," presented at Katalistiks' 4th annual FCC Symposium, Amsterdam, The Netherlands, May 18-19, 1983. 11. Peeples, J. E., "The Clean Air Act, a Brave New World for Fuel Reformulation," Fuel Reformulation, Vol. 3, No. 6, November/December 1993. 2. Dharia, D., Brahn, M., and Letzsch, W., "Technologies for Reducing FCC Emissions," presented at Stone & Webster's 5th annual Refining Seminar, Dallas, Texas October 12, 1983. 3. Yergin, D. and Lindemer, K., "Refining Industry's Future," Fuel Reformulation, Vol. 3, No. 4, July/August 1993. 4. Perino, J. O., "Blending Control Upgrade Projects," Fuel Reformulation, Vol. 3, No. 4, July/August, 1993. 5. Clarke, R. H. and Ritz, G. P., "Method for the Analysis of Complex Mix of Oxygenates in Transportation Fuels," Fuel Reformulation, Vol. 3, No. 4, July/August, 1993.
Emerging Trends in Fluidized Catalytic Cracking
33?
6. Urizelman, G. H., "NOX," Fuel Reformulation, Vol. 1, No. 6, November/ December 1991. 7. Piel, W. J., and Thomas, R. X., "Oxygenates for Reformulated Gasoline," Hydrocarbon Processing, July 1990, pp. 68-73, 8. Hirshfeld, D. S. and Kolb, J., "Minimize the Cost of Producing Reformulated Gasoline," Fuel Reformulation, Vol. 4, No, 2 March/April 1994 9. Unzelman, G. H., "A Sticky Point for Refiners," Fuel Reformulation, Vol. 2, No. 4, July/August 1992. 0. Nocca, J. L., Forestiere, A., and Cosyns, J., "Diversify Process Strategies for Reformulated Gasoline," Fuel Reformulation, Vol. 4, No. 4, September/ October 1994. 1. Desai, P. H., Lee, S. L., Jonker, R. J., De Boer, ML, Vending, J., and Sarli, M. S., "Reduce Sulfur in FCC Gasoline," Fuel Reformulation, Vol. 4, No. 6, November/December 1994. 2. Sarathy, P. R., "Profit from Refinery Olefins," Fuel Reformulation, Vol. 3, No. 5, September/October 1993. 3. Hosteller, R. and Cain, M., BP Oil, private communication, 1995. 4. Reid, T. A., Akzo Nobel, private communication, 1995. 5. Swain, E. J., "U.S. Crude Slate Continues to Get Heavier, Higher in Sulfur," Oil & Gas Journal, January 9, 1995, pp. 37–42. 6. Dougan, T. J., Alkemade, V, Lakhampel, B., and Brock, L. T., "Advances in FCC Vanadium Tolerance," NPRA Annual Meeting, San Antonio, Texas, March 20, 1994, reprinted in Grace Davison Catalagram. 7. Cunic, J. D., Diener, R., and Ellis, E. G., Exxon Research and Engineering, "Scrubbing—Best Demonstrated Technology for FCC Emission Control," presented at NPRA Annual Meeting, San Antonio, Texas, 1990.
APPENDIX 1
Temperature Variation
of Liquid Vis
^HnnHHIHIIIHHIiniUIIIIIIMIHIIIIIIIIIIIIIIIIMIIIIIIIIIIIIIIIIHIIIIIIIIIIIIItllllilllilllUlllllilllllllllllllllMllfllliWIIItlll MnnitMIIIIIIIIIIIIIIIHIMIIIIIIIIIIItlllllllllllllllllllllllllllllllllllllillHieHlllllllillllltllllliSKUMIitHIUitilMfllHIIilllll
Source: U.S. Department of Commerce, adapted from ASTM D-342-39.
338
APPENDIX 2
Correction to Volumetric Average Boiling Point WABP C 80) F V
==• ""^ WABP O 60 3 F VABP
A8TM Diet, 10% - 90 % Slop*
339
APPENDIX 3
TOTAL Correlations
Aromatic Carbon Content: CA = -814.136 + 635.192 x RI(20) - 129.266 x SG + 0.1013 x MW - 0.340 x S - 6.872 x ln(v)
Hydrogen Content: H2 = 52.825 - 14.26 x RI(20) - 21.329 x SG - 0.0024 x MW - 0.052 x S + 0.757 x ln(v)
Molecular Weight: MW = 7.8312 x 10-3 x SG-0-0976 x AP°C1238
Refractive Index @ 20°C: RI(20) = 1 + 0.8447 x SG1-2056 x (VABPoc+273.16r)0557 x
Refractive Index @ 60°C: RK60) = 1 + 0.8156 x SG12392 x (VABP0(: + 273.16)-0.0576 x
ource: Dhulesia, H., "New Correlations Predict FCC Feed Characterization Paramters," Oil & Gas Journal, Jan. 13, 1986, pp. 51-54.
340
APPENDIX 4
n-d-M Correlations v = 2.5 x (RI20OC - 1.4750) - (d2()OC - 0.8510) 05 = (d2()OC - 0.8510) - 1.11 x (RI2fn, - 1.4750) If v is positive: %CA = 430 x v +
If v is negative: %CA = 670 x v +
3660
M
If 03 is positive: %CR = 8 2 0 x G J - 3 x S + 10,000/M 10,600 If 03 is negative: %CR = 1440 x 03 - 3S +
M
%CN = %CR — %CA %C_r = 100— %CRK
Average Number of Aromatic Rings per Molecule (RA): RA = 0.44 + 0.055 x M x v
If v is positive
R^ = 0.44 + 0.080 x M x v
If v is negative
Average Total Number of Rings per Molecule (RT): RT = 1.33 + 0.146 x M x (03 - 0.005 x S)
If 03 is positive
RN = RT — RA
RT = 1.33 + 0.180 x M x (03 - 0.005 x S)
If 05 is negative
Average Number of Napthene Rings per Molecule (RN): R
M
=
RT—RA
ource: ASTM Standard D-3238-80. Copyright ASTM. Used with permission.
341
APPENDIX 5
Estimation of Molecular Weight of Petroleum Ofts from Viscosity Measurements
40 50 60 70 80 90 00 10 20 30 40 50 60 70 80 90
Tabulation of H Function H
334 355 372 386 398 408 416 424 431 437 443 448 453 457 461 465
336 357 374 387 399 409 417 425 432 438 443 449 453 458 462 466
339 359 375 388 400 410 418 425 432 438 444 449 454 458 462 466
341 361 377 390 401 410 419 426 433 439 444 450 454 459 463 466
343 363 378 391 402 411 420 427 433 439 445 450 455 459 463 467
342
345 364 380 392 403 412 420 428 434 440 446 450 455 460 463 467
347 366 381 393 404 413 421 428 435 441 446 451 456 460 464 468
349 368 382 394 405 414 422 429 435 441 447 451 456 460 464 468
352 369 384 395 406 415 423 430 436 442 447 452 456 461 465 468
354 371 385 397 407 415 423 430 437 442 448 452 457 461 465 469
Molecular Weight of Petroleum Oils Viscosity-Molecular Weight Chart LINES OF CONSTANT 210*F (98,89*C) VISCOSITY, cST
500
5
400
300
too
)0
j/
400
500
600
RELATIVE MOLECULAR MASS
Source: ASTM Standard D-2502-92. Copyright ASTM. Used with permission.
343
APPENDIX 6
Kinematic Viscosity to Saybolt Universal Viscosity Equivalent Saybolt Universal Viscosity, Sus
inematic Viscosity, cSt
1.81 2.71 4.26 7.37 10.33 13.08 15.66 18.12 20.54 43.0 64.6 86.2 108.0 129.5 139.8 151.0 172.6 194.2 215.8
At 100°F
At 210°F
32.0 35.0 40.0 50.0 60.0 70.0 80.0 90.0 100.0 200.0 300.0 400.0 500.0 600.0 648.0 700.0 800.0 900.0 1000.0
32.2 35.2 40.3 50.3 60.4 70.5 80.5 90.6 100.7 202.0 302.0 402.0 504.0 604.0 652.0
xtracted from ASTM Method D-2161-87. Copyright ASTM. Used with permission.
344
APPENDIX 7
API Correlations Xr = a + b x (R.) + c x (VG) Xn = d + e x (R.) + f x (VG) Xn = g + h x (R.) + i x (VG)
Where constants vary with molecular weight range given below:
Constants a b c d e f g h j
Heavy Fractions 200 < MW < 600 +2.5737 +1.0133 -3.573 +2.464 -3.6701 +1.96312 -4.0377 +2.6568 +1.60988
. = Refractivity Intercept VGC = Viscosity Gravity Constant R, K
K ~-R i(20)
Where:
j(2())
'
= Refractive Index @ 20°C = Density @ 20°C
ource: Riazi, M. R., and Daubert, T. E., "Prediction of the Composition of Petroleum ractions," Ind. Eng. Chem. Process Dev., Vol. 19, No. 2, 1982, pp. 289-294.
345
46
Fluid Catalytic Cracking Handbook
VGC = SG ~ °-24 - °-022 x log(V210 ~ 35.5) 0.755
Where:
V = Say bolt Universal Viscosity @ 210°F in seconds
Refractive Index @ 20°C (68°F):
I = A x exp(B x MeABP + C x SG + D x MeABP x SG) x MeABPE x SGF
Constants A B C D E F
2.341 * 10~2 6.464 x IQ"4 5.144 -3.289 x 10-4 -0.407 -3.333
MW = a x exp(b x MeABP + c x SG + d x MeABP x SG) x MeABP6 x SGf
Where:
Constants a b c d e f
20.486 1.165 x 10~4 -7.787 1.1582 x 10-3 1.26807 4.98308
APPENDIX 8
Definitions of Fluidization Terms
Aeration. Any supplemental gas (air, steam, nitrogen, etc.) that increases fluidity of the catalyst. Angle of Internal Friction—a. Angle of internal friction, or angle of shear, is the angle of solid against solid. It is the angle at which a catalyst will flow on itself in the nonfluidized state. For an FCC catalyst, this is about 80°. Angle of Repose—p. The angle that the slope of a poured catalyst will make with the horizontal. For an FCC catalyst, this is typically 30°.
SoHdSurfaca
347
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Fluid Catalytic Cracking Handbook
Apparent Bulk Density—ABD. The density of the catalyst at which it is shipped either in bulk volume or bags. It is density of the catalyst at minimum fluidization velocity. ed Density—pb. The average density of a fluidized bed of solid particles and gas. Bed density is mainly a function of gas velocity and, to a lesser extent, the temperature. Minimum Bubbling Velocity (Umb). The velocity at which discrete bubbles begin to form. Typical minimum bubbling velocity for an FCC catalyst is 0.03 ft/sec. Minimum Fluidization Velocity (Umf). The lowest velocity at which the full weight of catalyst is supported by the fluidization gas. It is the minimum gas velocity at which a packed bed of solid particles will begin to expand and behave as a fluid. For an FCC catalyst, the minimum fluidization velocity is about 0.02 ft/sec. article Density—p . The actual density of the solid particles taking into account any volume due to voids (pores) within the structure of the solid particles. Particle density is calculated as follows:
Po =
Skeletal density (Skeletal density x PV) + 1
ore Volume—PV. The volume of pores or voids in the catalyst particles. Ratio of Minimum Bubbling Velocity to Minimum Fluidization Velocity (Umb/Umf). This ratio can be calculated as follows: Umb
=
2300 x p°' 2< x n°"3 x exp° 7lteF
Umf= Where: pg = ji = F = dp = p = g =
d°8xg
