,+,+, applied catalysis A

ELS EV ] ER Applied Catalysis A: General 119 (1994) 121-138

Selective hydroisomerization of long chain normal paraffins

R o b e r t J. Tay lo r *, Randa l l H. Pe t ty Texaco. Inc., Research and Development Department, P.O. Box 1608, 4545 Savannah Ave., Port Arthur,

TX 77641, USA

Received 25 April 1994; revised 22 July 1994; accepted 26 July 1994

Abstract

The removal of long chain normal paraffins from lubricating oils is essential to produce an oil with acceptable cold flow properties. Conventional lubricant dewaxing processes remove these normal paraffins by selective cracking or crystallization mechanisms, resulting in a yield loss directly pro- portional to the concentration of the normal paraffins in the oil. More recent lubricant dewaxing processes attempt to isomerize the normal paraffins to isoparaffins, allowing them to remain in the oil and produce a higher quality oil in higher yield. The work reported here is a study of the effect of zeolite type and acidity on the selective isomerization of the long chain normal paraffins typically found in lubricating oils. Catalysts containing beta, USY, SDUSY, mordenite, ZSM-5 and SAPO-11 are tested for their ability to isomerize neat hexadecane and a 50:50 mixture of hexadecane and tetramethylpentadecane. Beta, SAPO-11 and ZSM-5 are also tested with a refined wax distillate lubricating oil which contains 15% normal paraffins. In summary, these results show that SAPO-11 was the only catalyst tested which was capable of isomerizing normal paraffins in the presence of isoparaffins without large yield losses due to unwanted cracking.

Keyworda: Dewaxing; Hydroisomerization; Lubricant oil; Paraffins (n-); SAPO-11

1. Introduction

1.1. Catalytic dewaxing

The discovery of ZSM-5 and its ability to discriminate between reactant mole- cules based on size and shape differences has had a major impact on chemical and petroleum catalytic processes. Mobil's catalytic dewaxing processes for distillates and lubricants are good examples of such shape-selective processes [ 1 ]. Catalytic

* Corresponding author. Fax. ( + 1-409) 9892826.

0926-860X/94/~.,07.00 © 1994 Elsevier Science B.V. All rights reserved SSDIO926-860X(94) 00 168-5

122 Robert J. Taylor, Randall H. Petty/Applied Catalysis A: General 119 (1994) 121-138

dewaxing removes paraffin type waxes by shape-selective hydrocracking, where the wax is converted to lower molecular weight products which are removed from the higher boiling lube oil by distillation. The ZSM-5 zeolite used in this catalyst has a pore opening which only admits the straight chain, waxy normal paraffins either alone or with slightly branched paraffins. More highly branched paraffin molecules, cycloaliphatics and larger aromatics are excluded from the pores. This results in a process that is very effective at removing normal paraffins and normal paraffin-like side chains. Removal of the normal paraffins is required to lower the pour point of the oil and improve its cold flow properties. Normal paraffins are usually to more than 95% removed [2].

With such extensive normal paraffin removal during catalytic dewaxing, it can be seen that as the normal paraffin content of a feed increases, the yield of dewaxed product will decrease. An improved process for dewaxing light, high normal par- affin content feedstocks is clearly needed. A more attractive dewaxing procedure would result from catalytic dewaxing through isomerization of the n-paraffins to multibranched isoparaffins, eliminating the yield loss associated with n-paraffin removal by cracking or solvent extraction [ 3]. While in principle this appears to be an acceptable method of dewaxing, in practice one finds that the wax present in many conventional feedstocks is not only composed of n-paraffins but also of isoparaffins, naphthenes and aromatics [4,5 ]. Even if complete isomerization of the n-paraffins could be obtained, acceptable pour points still may not be reached. The recent announcement of a new process for dewaxing hydrocracked lube stocks does illustrate the advantages of increased isomerization during the dewaxing process. [ 6-8 ] This process is reported to maintain a higher paraffinicity of the dewaxed oil relative to conventional catalytic dewaxing processes. This increased paraffinicity and lower cracking results because a portion of the wax is isomerized rather than cracked. This increased isomerization results in by-products which are not light gases and naphtha, as seen in conventional catalytic dewaxing, but low pour middle distillates and lower boiling lube range oils. While this new process does have some unique advantages over conventional catalytic dewaxing, its ability to dewax conventional solvent refined feedstocks, which make up 92% of the worldwide base oil production, is still uncertain. Difficulty in dewaxing solvent extracted feedstocks may result from the presence of non-paraffin wax components which are not found in the hydrocracked lube feeds used in the development of this new process.

Another approach to obtaining an improved dewaxing process for these high n- paraffin content feeds is to combine the isomerization and selective dewaxing steps into one process. [ 9,10] Synergistic effects from the combination of isomerization and cracking catalysts which give enhanced yields of multibranched isoparaffins in the bifunctional conversion of n-paraffins have been reported. [ 3 ]. Theoretically, such a combined process may be obtained by first hydroisomerizing the waxy feedstock prior to the selective dewaxing step. The objective of the isomerization step is to isomerize the normal paraffins in the feed to less waxy isoparaffins without

Robert J. Taylor, Randall H. Petty/Applied Catalysis A: General 119 (1994) 121-138 123

producing products outside the lube oil range. The severity of this step is controlled to carry out only a partial removal of the waxy components. Final dewaxing to the desired pour point is carried out during the selective dewaxing step, where the remaining waxy paraffin components are removed, leaving the newly formed branched isoparaffins in the product. The combination of these two steps may increase the yield across the dewaxing process, providing the isoparaffins formed are highly enough branched to remain in the dewaxed oil during the subsequent selective cracking step. Recent patent literature has reported the advantages of such a process using stacked beds of isomerization and cracking catalysts [ 11 ].

1.2. Selective hydroisomerization

The use of zeolite containing catalysts in hydrocarbon processing is extensive, including such processes as catalytic cracking, hydrocracking and hydroisomeri- zation. Catalytic cracking catalysts are monofunctional and take advantage of the acidity characteristics of the zeolite component to carry out the desired cracking reactions. Hydrocracking catalysts, on the other hand, are bifunctional and have both a zeolite component and an active metal component which give the catalysts both cracking and hydrogenation activity, respectively. When these bifunctional catalysts are used in hydroprocesses, the composition of the reaction products is dependent on the reaction temperature and the relative strength of the hydrogenation and cracking activities in the catalyst. The accepted reaction scheme for hydro- cracking normal paraffins is seen in Fig. 1 [ 12]. The first step in this scheme is the formation of olefins at the metallic sites, which is followed by the formation of carbenium ions at the acidic site. These ions may then isomerize to secondary or tertiary molecules or crack to smaller molecules, depending on the reaction con- ditions and the nature of the catalyst present. Olefins formed in these reactions are in turn saturated at the metal site to give mostly a paraffinic product. Hydroiso- merization is differentiated from hydrocracking by the degree of isomerization versus cracking which takes place. Typically, hydroisomerization dominates at less severe reaction conditions and low degrees of overall conversion.

The acidity of the catalyst has a major influence on the hydroisomerization and hydrocracking mechanisms. The acid site density and acid strength distribution are both important and the proper balance of these variables is critical in determining the reactivity and selectivity of these catalysts. Catalysts which have a high degree of hydrogenation activity and a low degree of acidity are best for maximizing hydroisomerization versus hydrocracking. The pore opening of the molecular sieve can also have a major effect on the selectivity of these catalysts. Table 1 lists the size of the pore openings for the molecular sieves tested here. If the pore opening is small enough to restrict the larger isoparaffins from reacting at the acidic sites inside the pore, the catalyst will show good selectivity for converting normal paraffins. The ideal catalyst for the selective hydroisomerization of n-paraffins should have both selectivity for isomerization, which comes from the proper balance

124 Robert J. Taylor, Randall H. Petty / Apphed Catalysis A: General 119 (1994) 121-138

HYDROGENOLYSlS n-ALKANE - - ~ CRACKED

PRODUCTS

T-2 H +

÷H

n-ALKENES ~ SEC. n-ALKYL CATIONS 4 - -

i-ALKENES

+H +

CRACKED PRODUCTS

REARRANGEMENT

TERT. i-ALKYL CATIONS ~ CRACKED PRODUCTS

!+2H

i -ALKANES HYDROGENOLYSIS ~ CRACKED

PRODUCTS

Fig. 1. Reaction scheme for hydrocracking of normal paraffins [ 12].

Table 1 Zeolite structure information [ 13]

Zeolite name Zeohte type index Pore opening angstroms

ZSM-5 MFI 5.3 × 5.6 5.1 ×5.5

SAPO- 11 AEL 3.9 × 6.3 Beta [ 14] BEA 5.7 × 7.5

5.6X6.5 Mordenite MOR 6.5 x 7.0 USY FAU 7.4 × 7.4 SDUSY FAU 7.4 × 7.4

of acidity and hydrogenation activity, and selectivity for reacting only with normal paraffins, which comes from the size of the pore opening of the molecular sieve used.

2. Experimental

2.1. Catalyst preparation procedures

Catalyst supports were obtained from a variety of catalyst vendors. The ultrast- able zeolite Y (USY), super dealuminated USY (SDUSY), and beta catalysts were

Robert J. Taylor, Randall 14. Pet~ / Applied Catalysis A: General 119 (1994) 121-138 125

obtained from the PQ Corporation (Valfor CP316-56, CP316-26 and C861B). The mordenite catalyst was obtained from Engelhard (EZ-320E). The ZSM-5 and SAPO-11 catalysts were obtained from UOP (MFI and B-287). These supports were obtained from the catalyst vendors as ~6 inch extrudates prepared with zeolite and an alumina binder in an 80/20 ratio. The 50% beta support was a ~6 inch extrudate prepared using a 50/50 weight percent mix of beta zeolite powder from PQ Corporation and alumina from Vista.

The extrudates were ground and sieved using No. 20 and No. 30 trays. The 20- 30 material was collected, dried, calcined and used as the catalyst support for this study. This support was then impregnated with an aqueous solution of palladium tetraammine chloride using incipient wetness procedures. Enough palladium was used to prepare a 0.5 wt.-% Pd catalyst. Finished catalysts were then dried and calcined before use.

2.2. Catalyst acidity measurements

The catalyst supports were analyzed using temperature-programmed desorption of chemisorbed ammonia. This test yields information concerning both the total acid site density and acid strength distribution of the material. All the catalyst supports were analyzed using the same procedure. The values of total acidity are reported in units of milliequivalents of ammonia per gram of sample and are accurate for comparison of data collected using the same apparatus and procedure. The values are reproducible to within 2-3%. The objective of this test is to provide an internally consistent ranking of catalysts by relative acidity and not an absolute measure of the total number of acid sites.

Powdered samples were allowed to equilibrate with atmospheric moisture prior to testing. The weight loss at 538°C was then determined to correct the sample weight for any water present. Samples of 100 to 130 mg were housed in a quartz tubular reactor (6 mm I.D.) and pretreated in flowing helium (20 sccm) while heating at 10°C/rain to 538°C. Samples were allowed to remain at 538°C overnight. The samples were then cooled to 175°C and ammonia was adsorbed onto the sample. Typically, a 5 tool-% ammonia in helium mixture (Scott Specialty Gases) was passed over the sample for 30 min to saturate the surface with ammonia. The effluent from the reactor bed was tested during the adsorption step to be sure surface saturation was obtained. The sample was then allowed to equilibrate in flowing helium (20 sccm) for 1 h at 175°C. Following the equilibration, the temperature was dropped to 75°C and the ammonia was desorbed from the surface using a linear temperature ramp of 10°C/rain to 650°C.

The outputs from the temperature controller and gas detector were collected using a computer equipped with an I /O board and data acquisition software. The area under the curve was integrated from 75°C to 500°C and this value was used to determine the total acid site density of the catalyst. In order to determine the acid strength distribution, the desorption profiles were fit using three peaks and the

126 Robert J. Taylor, Randall H. Petty /Applied Catalysls A: General 119(1994) 121-138

maxima and widths of the peaks were held as constant as possible while fitting each profile. The r-squared for each of the fits was greater than 0.995. The total area and the area under each of the peaks were determined and the normalized percentage was calculated for each peak. Weak, medium and strong acidities are defined as the areas under the peaks at the lowest, medium and highest temperatures. The use of three peaks to fit the profiles was not based on peak assignment to any specific acid site or type (BrCnsted or Lewis) but, instead, was a convenient method to categorize the acid strength distributions obtained by this method.

2.3. Catalyst screening procedures

Model compound experiments n-Hexadecane and a 50:50 weight percent mixture of n-hexadecane (n-C16) and

2,6,10,14-tetramethylpentadecane ( i-C 19) were used as feedstocks in the se exper- iments. The experiments were carried out in a small fixed bed reactor in down-flow configuration. The reactor had a 0.37 inch internal diameter (! inch SS tubing) 2

Glass beads were used to support the catalyst bed in the reactor tube. The reactor was loaded with 10 ml of catalyst (nominal 20-30 mesh particles). The weight of catalyst varied depending on its density. Six inches of glass beads above the catalyst bed served as a preheat zone for the feed in the reactor. Two thermocouples were placed on the exterior of the reactor at the top and the bottom of the catalyst bed. The temperature was controlled to ___ 2°F ( l°F = 0.55°C). The catalysts were activated by gradual heating to 400°F (204°C) (overa4hper iod) under a hydrogen flow of 330 cm3/min at 500 psig ( 1 psig = 6.895 kPa). After a 1 h hold at 400°F, the catalyst was cooled to 350°F and the feed introduced to the reactor at a liquid hourly space velocity (LHSV) of 1 h - 1.

The conditions were allowed to stabilize over a 12 h period at 300°F. Samples were collected at 4 h intervals. The liquid product was then analyzed by gas chromatography for unreacted feed, isomerized and cracked products. The total liquid product yield was calculated by dividing the weight of liquid product col- lected by the weight of feed entering the reactor during the collection period. The gas yield was calculated as the difference between 100 and the percent liquid yield. The gas was not analyzed.

Wax distillate experiments The catalyst screening runs using the refined wax distillate oil were made in a 1

liter continuous stirred tank reactor with a Berty reactor design. In each run, 65 cm 3

of catalyst were placed in the catalyst basket inside the reactor. The reactor had a measured liquid holdup volume of 125 ml. The reactor was heated under 1000 psig hydrogen flowing at 210 sccm from ambient temperature to 450°F over 10 h. The reactor temperature was then adjusted to the desired start of run temperature and the reactor liquid and gas feed rates were adjusted downward over 4 h from 300 cm3/h to 50 cm3/h and 1250 sccm to 210 sccm, respectively. The impeller was set

Robert J. Taylor, Randall H. Petty/Applied Catalysis A: General 119 (1994) 121-138 127

to 1250 rpm. After a 4 h equilibration period at constant conditions, the product was collected every 4 h. For each sample collected, the product weight and volume were measured and the total liquid product yield was calculated based on the weight entering the reactor and the amount collected. Typically, the reactor was held at constant temperature for 16 h and the four samples collected during that period were composited, yielding approximately one quart of product. The temperature was raised over the course of the run and a 4 h transition sample was taken each time the temperature was raised. This sample was not included in the composites.

Each of the composite samples were initially tested for pour point and boiling range. If the boiling range showed no light ends, then the product was used as the hydrorefined (HR) oil. However, if the boiling range showed greater than 5% boiling less than 610°F, then the composites were fractionated using an ASTM D- 1160 type distillation. The procedure was varied from the D-1160 method by using a larger volume and higher vacuum. The 610 + °F fraction from this fractionation was used as the HR oil. The HR oils were tested for boiling range, pour point, wax content, and n-paraffin content. The HR oil yield was calculated by correcting the total liquid product yield for the 6 1 0 - °F fraction in the total liquid product.

In order to compare the results obtained from catalytic dewaxing and solvent dewaxing, the refined wax distillate oil was solvent dewaxed. The solvent dewaxing experiments were carried out using a pilot scale, batch dewaxing apparatus and the conditions were chosen to simulate commercial operations. The feedstock was diluted 5:1 in a 50:50 MEK/toluene solvent and heated to its complete miscibility point. It was then cooled at 5°F per minute to 15°F below the desired dewaxed oil pour point. The wax was then separated by filtration and the wax cake was washed with solvent at the filtration temperature in a 5:1 ratio of solvent to feedstock. The solvent was then stripped out of the dewaxed oil in a batch still.

The wax contents of the feedstock and products were calculated using a differ- ential scanning calorimetry method described elsewhere [ 15 ]. This method deter- mines the dry wax content at any specific pour point and has been shown to be as precise and accurate as any existing method, including physical separation methods.

The normal paraffin content of the hydrorefined oils was determined with a Carlo-Erba high-temperature gas chromatography (GC) method. The sample was initially analyzed without treatment. It was then stripped of normal paraffins by refluxing with 5A molecular sieves in toluene and the resulting sample again analyzed by high-temperature GC. The n-paraffin content was determined by dif- ference. Errors due to volatility losses were corrected for by referencing to an internal standard. Results are reported on a weight percent basis.

3. Results

3.1. Acidity measurements

Fig. 2 shows the ammonia temperature-programmed desorption profile for each of the catalyst supports. It can be seen that these catalysts show a wide range of

28 Robert J. Taylor, Randall H. Petty/Applied Catalysis A: General 119 (1994) 121-138

S A P O - 1 1

i iIill i f

U S Y

, I , I i I , , I i I , I

. . . . MonoENirE " ~ " . . . . . . . . . . . . . SOusY . . . . . . . . . .

. . . . . . . . . . //_° _ _ _ ' ~ . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

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B E T A Z S M - 5

i k.-V., \ ,S . . . "% % • % . _ . . _ j , , - . . . ' , . " . . , - - . . .

200 400 600 0 200 400 600

Temperature, C Temperature, C

Fig. 2. Ammonia temperature-programmed desorption profiles for catalyst supports.

Robert J. Taylor, Randall H. Petty~Applied Catalysts A: General 119 (1994) 121-138

Table 2 Total aodity and acid strength distribution of catalyst supports

129

Catal yst Total acidity Acid strength distribution (meq/g)

Weak acidity Medium acidity Strong acidity

Center C Width C % of Center C Width C % of Cent er Width C % of Total Total C Total

SAPO-I1 0.27 269 43 34 345 58 52 442 62 14 Beta 0.71 269 43 34 355 57 42 442 62 24 Morde mte 0 94 276 48 26 387 52 8 539 99 65 SDUSY 0.22 280 41 34 370 49 39 452 65 27 USY 0.85 277 41 33 355 49 28 452 73 39 ZSM-5 0.71 269 43 23 355 57 20 450 62 57

acid site density and acid strength distributions. Table 2 gives a quantitative com- parison of these values, where the total acidity is the total amount of ammonia desorbed by the support and the acid strength distribution is the normalized area under each of the three guassian peaks used to fit the desorption profile. The peak maxima and widths are also listed for the peaks used to fit the desorption curve. It can be seen in Table 2 that each of the catalyst supports has a similar percentage of weak acid sites but there is a wide range of percentages in the medium and strong sites.

3.2. Hydroisomerization of n-C16

The catalysts were first tested for their ability to isomerize n-hexadecane. Fig. 3 shows a plot of the product distribution (isomerized and cracked), liquid yield and n-hexadecane conversion for four of the catalysts tested. The results for the SDUSY and beta catalysts show that as the temperature is increased, the initial reaction which occurs is the isomerization of the n-hexadecane to C16 isoparaffins. As the temperature is further increased, the amount of isomerized product decreases and the amount of lighter cracked products increases. At the highest temperature tested the liquid yield began to drop significantly, indicating cracking to gas products. The graphs for ZSM-5 and SAPO-11 show significantly different results. In the case of ZSM-5, no significant isomerization activity is seen with the major product being species from cracking reactions. The SAPO-11 catalyst gives a high degree of isomerization, with 70% isomerized product at 97% conversion. Selectivities for n-hexadecane isomerization of 85% at 94% conversion have been reported using Pt/SAPO-11 [ 16]. It can also be seen that the SAPO-11 catalyst requires a significantly higher temperature to carry out the isomerization reaction. This decrease in activity can be explained by the lower total acidity and acid strength of the SAPO-I 1 support.

Fig. 4 shows the a plot of the percent of isomerized product formed versus the percent of total conversion of n-Ct6. It can be seen in this graph that all of the large

130 Robert J. Taylor, Randall H. Petty/Applied Catalysis A." General 119 (1994) 121-138

100

SDUSY SAPO-11

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40 W W n'-

W O

20

300 350 ~ 0 450 500

REACTOR TEMPERATURE, F

100

8O

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0 300 350 400 450 500 550 600

REACTOR TEMPERATURE, F

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100 100 ~ - I

80 ~ 8o . . . . . . . . . . . . . . . . / - - \ . . . . . /

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o , . o . . . . 3o0 350 400 480 3o0 3so 4oo 460 5o0

REACTOR TEMPERATURE, F REACTOR TEMPERATURE, F

Fig. 3. Plot of the product distribution, yield and conversion versus reactor temperature for the reaction of n- hexadecane with SDUSY, SAPO-I 1, ZSM-5 and beta catalysts. (O) Liquid yield; (11) conversion; ( * ) Isom- enzed product, ( * ) Cracked product in liquid.

pore zeolites (beta, mordenite, SDUSY, and USY) are capable of carrying out some degree of isomerization. The high activity of the USY sample resulted in a deviation from the others because the conversion which gives the maximum isom- erization was not obtained. It is clear from the graph that of the two intermediate pore size catalysts, only the SAPO-11 carried out any significant isomerization. The ZSM-5 showed less that 10% isomerized product formed.

These results are in line with the proposed mechanism for hydroisomerization of n-alkanes by bifunctional catalysts shown in Fig. 1. At low conversions, in this case less than 40% total conversion of n-C16, isomerization dominates in all but the

Robert J. Taylor, Randall H. Petty /Applied Catalysis A: General 119(1994) 121-138

80

131

60 . . . . . . . . . . . . . . . . . . . . . . . . . . . . .:..-"-; , . 4 . . . ~ .....

40 . . . . . . . . . . . . . . . . . . ~ - - - - - ' . , . - . - - : . . . . .

o

. . . . . . . . . . . . . - .

0 20 40 60 80 100 TOTAL CONVERSION (%)

Fig. 4. Plot of the percent of isomerized C~6 product formed versus total conversion of n-hexadecane. ( - - I I - - ) SAPO-11; ( - ' - 0 - . - ) Beta; ( - • - ) Mordenite; ( - - * - - ) SDUSY; ( . . . . . . [ ] . . . . . . ) USY; ( - O - ) ZSM-5.

ZSM-5 sample. As conversion increases, cracking increases through both direct cracking of the carbocation intermediates and secondary cracking of the isoparaf- fins. In the case of ZSM-5, the cracking reactions dominate at all conversion levels tested.

3.3. Hydroisomerization of n-C ]o/i-C19

Reactant selectivity Fig. 5 shows a plot of the percent of conversion of n-Cl6 versus the total con-

version for the reaction of each of the catalysts with the i-C19/n-C16 mixture. This graph shows that only the intermediate pore size catalysts have a good selectivity for reacting with the n-C16 in the mixed feed. This comes from their restrictive pore size. All of the catalysts containing larger pore zeolites show low activity for converting the n-C16 until nearly all of the i-C19 has been converted. This is a clear example of reactant selectivity based the size of the pore opening. The large pore catalysts show no selectivity based on molecule size but instead convert the more reactive i-C19 first. In the case of the intermediate pore size catalysts, the less bulky n-C~6 was converted first instead of the more reactive i-C19. For ZSM-5, less than 10% of the i-Cl9 was converted under the conditions tested.

Hydroisomerization selectivity Fig. 6 shows a plot of the product distribution, as isomerized C16 , isomerized

C19 and cracked product, and the liquid yield versus the reactor temperature for four of the catalysts tested with the i-C]9/n-C]6 mixture. Again, the SDUSY and

132 Robert J. Taylor, Randall H. Petty/Apphed Catalysis A." General 119 (1994) 121-138

100 . . . . . . . . . . . . . . . . . . . . -./5, . . . . . . . . . . . . . .

I.J.I

~/- J W > 60 ." Z O jo t~ t" ,¢,D •

40 C •

20 . . . . -Z; " 7~"¢ . . . . . . . . . . . . . . . . . . .

~ o o°

0 I i

0 20 40 60 80 100 % TOTAL CONVERSION

Fig. 5. Plot of the percent of n-Cl6 converted from a 50:50 mixture of n-C16/i-C29 versus total conversion. ( - - I I - - ) SAPO-11; ( - - - 0 - . - ) Beta; ( - • - ) Mordemte; ( - - - - A , - - ) SDUSY; ( . . . . . . [] . . . . . . ) USY; ( - O - ) ZSM-5.

Beta catalysts give similar results, showing no shape selectivity. In both cases, the initial reaction is the isomerization of the i-C19. As the reaction temperature is increased, the isomerized C19 product begins to crack before any significant amount of the n-C16 begins to isomerize. In contrast, the SAPO-11 catalyst shows good activity for isomerizing the n-C16 as well as some i-C19. The ZSM-5 catalyst, which also showed some shape selectivity, only shows cracking activity at the conditions tested.

Fig. 7 shows the percent ofn-C 16 isomerized versus the percent of n-C16 converted and the percent i-Cj9 isomerized versus the percent i-Ct9 converted. It can be seen from these graphs that of the catalysts which have good reactant selectivity for n- C16, i.e. ZSM-5 and SAPO-11, only SAPO-11 also shows good hydroisomerization activity. While the ZSM-5 selectively converts n-C16 , it predominately produces cracked product. For the catalysts containing the larger pore zeolites, the extent of hydroisomerization activity seems to be affected most by the acidity. The higher the total acidity, the lower the hydroisomerization selectivity. The exception is beta zeolite, which shows both high acidity and good hydroisomerization selectivity.

Fig. 8 shows a plot of the percent of the feed converted to i -C16 v e r s u s the percent of feed converted to cracked product. It can be seen that while all the catalysts except ZSM-5 will isomerize n-C16, only SAPO-11 will do so at low cracking activities.

3.4. Hydroisomerization o f a refined wax distillate

Catalysts containing SAPO-11, beta and ZSM-5 were evaluated using a refined wax distillate feedstock containing 15 wt.-% normal paraffins. The properties of

Robert J Taylor, Randall H. Petty/Applied Catalyszs A: General 119 (1994) 121-138 133

100

8O

£3 O n" 60 13_ Q

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SDUSY SAPO-11

20

300 350 400 450

REACTOR TEMPERATURE, F

1 8 o 6° i 40

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0 - - I 400 450 500 550 600

REACTOR TEMPERATURE, F

100

8O

E~ E3 O C]: 6O O_ 0 D 0 ._1 4O Z

ZSM-5

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100

80

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. . . . . . . . . . . . . . . A Y

300 350 400 450

REACTOR TEMPERATURE, F

20

BETA

300 320 340 360 380

REACTOR TEMPERATURE, F Fig. 6 Plot of the product distribution (isomerized C~6, isomerized C~9, cracked products) and yield versus temperature for the reaction of a 50:50 mixture of n-Cl6/i-C19 with SDUSY, SAPO-11, ZSM-5 and beta catalysts. (O) Liquid yield; ( • ) isomerized n-C 16; ( * ) isomerized i-Cl9; (~r) cracked product in liquid.

this feedstock are shown in Table 3. The results from these evaluations are sum- marized in Fig. 9 along with results obtained from dewaxing the same feedstock using a solvent dewaxing process. The top graph in this figure shows a plot of the pour point of the hydrogen refined oil versus the yield loss obtained while producing the refined oil. The pour point of an oil is the temperature at which it is unable to flow. From the top graph in Fig. 9, it can be seen that the solvent dewaxing process has the best selectivity for lowering the pour point, giving the largest pour point reduction with the smallest yield loss. For the catalytic systems, the ZSM-5 and

134 Robert J. Taylor. Randall H. Petty/Applied Catalysts A: General 119 (1994) 121-138

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0 0

. . . . • .:" " ..... "-'-" ......

. . ' t . . . . . . t~J I".. , ....,, . . . . . . . . . . . . . . . . : ~ - ; . . . . . . . . . , , . , ' - ' s . . . . . . .~ . . . . . . . -": ' -"- . . . . . . . . . . . . \ - -~- il -

,..

m ~ . . o . . 4 > , , ,

20 40 60 80 100 % i-019 CONVERTED

Fig. 7. (Top) Plot of the percent of n-C16 isomerized versus the percent conversion of n-C~6. (Bottom) Plot of the percent of i-C~9 isomerized versus the percent conversion of i-C19. (__m__) SAPO-1 l; ( - . - • - . - ) Beta; ( - • - ) Morclenite; ( - - -k - - ) SDUSY; ( . . . . . . [] . . . . . . ) USY; ( - O - ) ZSM-5.

SAPO-11 catalysts give the best results with both the 50% and 80% beta catalysts showing the worst selectivity.

The middle graph in Fig. 9 shows a plot of the percent of wax removal versus yield loss, which shows the selectivity of the catalysts for removing wax compo- nents. It should be noted that removal of wax does not mean loss of yield. The

Robert J. Taylor Randall H. Pettv lApplied Catalysis A." General 119 (1994) 121-138 135

9 t- O LU

r r

O CO o Q LU F- rr LU > Z O

D LU LLI LL

40

30

20

10

0 0

• ," . I . . . . . . .~. .......... 41t ".. 2 . . . . . . . ,.,." ,,

• , ' " ~ ,o "°# "'~.. % • . . j . - . . . . . ~ . ":

I ," . S > * ..;." . . . . . . . . . . . . . . . . . . . . . -'.'~.~ \ . . . . . • ,a "/ r--~."" ~'.._~

. . . . . . . . .o ' " ~ " ~ :

....... ~ . " ~ ....................... ,---.,,- ........ ~'-\ , I ~ 1 ~

20 40 60 80 100 % FEED CONVERTED TO CRACKED PRODUCT

Fig. 8. Plot of the percent of feed converted to isomerized C16 versus the percent of feed converted to cracked products. ( - - I - - ) SAPO-I1; ( - - - 0 - . - ) Beta; ( - • - ) Mordenite; ( - - - ~ - - - ) SDUSY; ( . . . . . . [] . . . . . . ) USY, ( - ~ - ) ZSM-5.

objective of a wax isomerization step in a combined process is to convert the waxy components to non-wax components which remain in the oil during the second dewaxing step. For this to occur, the catalyst must selectively isomerize the wax to non-wax and must not crack oil molecules out of the oil fraction• In the case of solvent dewaxing, not only is there a yield loss from the removal of the wax but also from the removal of some oil which is trapped in the wax during crystallization. The results in the middle graph of Fig. 9 show that all the catalysts except the 80%

Table 3 Refined wax distillate feedstock properties

Pour point (°F) 80 Viscosity, 40°C (cSt) 17.267 Viscosity, 100°C (cSt) 3.808 Viscosity index 111 Viscosity, SUS, 100°F 92 n-Paraffin content ( wt.-% ) 15.3 Wax content (wt.-%) 20.7

SIMDIS(wt.-%) IBP/5 605/670 10/20 690/713 30/40 730/745 50/60 759/774 70/80 790/809 90/95 840/868 FBP 950

136 Robert J. Taylor Randall H. Petty/Apphed Catalysis A: General 119 (1994) 121-138

100 . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

so -- .~.~"-" -.- . . . . . . . . . . . . . . . . . . . . . . . . . .

n ° 80

" - . , . . . . . . '° . ., : - - , . . ~ .

20

0 P , ~ I t t , I ,

100 . . . . . . . . . . . ~f]~[ . . . . . . . . . . . . . . . . . . . . . . . . . . .

~> eo

~: 60

x I - ~ ,o . . . . . . - : ' - : ' - : ' - . . . . . . . . . . . . . . . . . . . .

~ - 20 . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

o ~ p , I r I , r , [ ~ I ,

100 ~

~ o

# . = o . ' q v

o p , - - ~ 7 , ; ? I I i I I

0 10 20 30 40 .60 60 70 HR Oil Yield Loss, wt%

Fig. 9. Plots of the hydrorefined oil pour point (top), percent of wax removal (middle) and percent of normal paraffin removal (bottom) versus the hydrorefined oil yield loss. ( - - I I I - - ) SAPO-] 1; (--- • - .-) Beta; ( - • - ) M o r d e n i t e ; ( - - - ~r - - - ) S D U S Y ; ( . . . . . . [ ] . . . . . . ) U S Y ; ( - ~ - ) Z S M - 5 .

Robert J. Taylor, Randall H. Petty~Applied Catalysis A" General 119 (1994) 121-138 137

Beta catalyst give similar results up to 50% wax removal. The most selective at high wax removals is solvent dewaxing. Again, both the SAPO-11 and ZSM-5 catalyst show similar selectivities; however, based on the previous model compound study, it is expected that the SAPO- 11 catalyst gives the highest degree of isom- erization. These results suggest, however, that any yield advantage obtained from isomerization of wax to oil with SAPO-11 is cancelled out by yield loss due to non- selective cracking.

The bottom graph in Fig. 9 shows a plot of the percent of n-paraffin removal versus the yield loss. Results here are similar to those observed with the wax selectivity. Solvent dewaxing and the ZSM-5 catalyst are the most selective for removing n-paraffins but are not expected to convert any n-paraffins to isoparaffins and increase the oil yield. SAPO-11 shows fairly good selectivity for removing n- paraffins up to about 60% removal but then begins to give significant yield loss. The loss in selectivity for the SAPO-11 catalyst may result from the higher tem- perature required to obtain the desired isomerization reactions. Unfortunately, the complexity of the feed did not allow the degree of isomerization of the n-paraffins to be determined. Based on model compound studies, it can be assumed that at least half of the n-paraffins removed using SAPO-11 were isomerized but the possible yield advantage from carrying out this isomerization is cancelled out by the non- selective cracking of the oil molecules.

4. Conclusions

Experimental results using model feedstocks of n-hexadecane have shown that normal paraffins can be isomerized without significant cracking in the absence of isoparaffins. However, when a mixture of n-hexadecane and tetramethylpentade- cane is used, it is difficult to isomerize the normal paraffins before the isoparaffins begin to crack. SAPO-11 was the only catalyst which was found capable of isom- erizing normal paraffins in the presence of isoparaffins without large yield losses due to cracking in these model compound studies. This catalyst has both the selectivity for isomerization, which comes from the proper balance of acidity and hydrogenation activity, and selectivity for reacting only normal paraffins, which comes from the size of the pore opening of the molecular sieve present. When the SAPO-11 catalyst was used to dewax a refined wax distillate feed, no yield advan- tage was seen relative to solvent dewaxing or ZSM-5 catalyst. It is suggested that any yield advantage obtained from isomerization using the SAPO-11 catalyst is offset by a yield loss from non-selective cracking of the oil molecules.

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138 Robert J. Taylor, Randall 11. Petty~Applied Catalysis A: General 119 (1994) 121-138

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