USY Zeolite Catalysts

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Article

VGO Hydrocracking on NiMo / US Y zeolite catalysts. Experimental study and kinetic modeling Tao Zhang, Carolina Leyva, Gilbert F. Froment, and Jorge Martinis Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/ie503567b • Publication Date (Web): 10 Dec 2014 Downloaded from http://pubs.acs.org on December 23, 2014

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Industrial & Engineering Chemistry Research

VGO Hydrocracking on NiMo / US Y zeolite catalysts. Experimental study and kinetic modeling Tao Zhang*, Carolina Leyva** and Gilbert F. Froment*** Texas A & M University, College Station 3122, Texas & Jorge Martinis, Bryan Research & Engineering, Bryan, Texas 77805

*Present address: China University of Petroleum, Beijing, 10224 ** Present address: Centro de Investigacion en Ciencia Applicada y Tecnologia Avancada-IPN-Mexico City 11500,Mexico ***Corresponding author; [email protected]

Abstract A bench-scale experimental unit based on a Robinson-Mahoney reactor with completely mixed gas and liquid phases was used to study the hydrocracking of a light vacuum gas oil on two base metal sulfide containing acid catalysts, characterized by their textural properties, NH3-TPD and pyridine-adsorbed Fourier-transformed infrared (Py-FTIR) acidity. The reactor effluent was analyzed in great detail by means of on-line GC and GC-MS and evidenced the role of the catalyst acidity. The detailed analysis allowed the reaction scheme to be expressed at the level required by the kinetic analysis in terms of the fundamental Single Event Kinetics approach and thus drastically reduce the number of independent kinetic parameters to be determined from the experimental data. Reactor simulations illustrate the detailed predictions made possible by this approach.

1. Objective The objective of this experimental program on the hydrocracking of Vacuum Gas Oil (VGO) was to produce detailed experimental data on the influence of space time, temperature and H2/VGO ratio on the conversion and yield patterns of the products for 2 typical hydrocracking catalysts with NiMo as metal components and USY zeolite with different acidity as carrier.

ACS Paragon Plus Environment


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For these catalysts the hydrocracking of a partially hydrogenated VGO was studied at temperatures in the range 360-395 °C, a pressure of 82 bar and molar ratios of hydrogen to VGO between 10 and 16. The VGO feed rate was varied to observe the evolution of the conversion and yield pattern at a given temperature and ratio of H2/VGO. The experiments were carried out in a Robinson-Mahoney fixed bed catalytic reactor with complete mixing of gas and liquid, also within the catalyst bed, i.e. operating under point conditions and thus eliminating problems of liquid and gas flow patterns in the model used to derive the conversions and yields from the experimental data. Each run required a detailed analytical effort on a TCD/FID GC and a GC-MS. The mass balance of each run was carefully checked. The detailed product analysis allowed the estimation of the parameters of the fundamental kinetic model of the hydrocracking process based upon the Single Event approach.

2. Feedstock analysis The feedstock is the product of the first stage of a hydrocracking unit with average molecular weight 233 g/mol. It may be labeled as light vacuum gas oil, or as heavy diesel. The Zdistribution of the compound types (or “classes”) as a function of the carbon number, determined by the NOISE method (Nitric Oxide Ionization Spectrometry Evaluation) is given in Table 1 and Figure 1.

C#\Z C5 C6 C7 C8 C9 C10 C11 C12 C13 C14 C15 C16 C17 C18 C19

+2 0.08 0.6 1.11 1.69 1.92 1.86 1.61 1.34 1.53 1.48 1.36 1.2 0.92 0.95 0.82

SATURATES 0 -2 0 0 0.5 0 1.95 0.03 3.24 0.12 4.58 0.71 3.74 1.39 2.92 2.54 2.23 3.34 2.12 4.08 1.84 3.45 1.63 2.78 1.47 2.07 1.2 1.58 1.09 1.36 0.85 1.07

-4 0 0 0 0 0 0.02 0.14 0.44 0.87 1.34 1.64 1.7 1.35 1.12 0.82

-6 0 0.01 0.14 0.41 0.5 0.45 0.34 0.33 0.39 0.42 0.46 0.48 0.47 0.44 0.37

MAR -8 0 0 0 0.02 0.13 0.32 0.49 0.53 0.47 0.41 0.36 0.33 0.31 0.22 0.15

DAR -10 0 0 0 0 0 0 0 0.03 0.06 0.15 0.29 0.35 0.31 0.24 0.15


-12 0 0 0 0 0 0 0.01 0.01 0.02 0.02 0.03 0.05 0.07 0.08 0.07

-14 0 0 0 0 0 0 0 0 0.01 0.03 0.05 0.06 0.05 0.04 0.03

TAR -18 0 0 0 0 0 0 0 0 0 0.01 0 0 0 0 0

Total 0.08 1.11 3.23 5.48 7.84 7.78 8.05 8.25 9.55 9.15 8.6 7.71 6.26 5.54 4.33

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C20 C21 C22 C23 C24 C25 Total

0.55 0.26 0.14 0.07 0.05 0.03

0.59 0.36 0.22 0.13 0.07 0.04

0.76 0.47 0.28 0.15 0.1 0.05

0.53 0.34 0.21 0.14 0.09 0.05

0.25 0.19 0.11 0.07 0.03 0.02

0.1 0.06 0.03 0.02 0.01 0

0.1 0.06 0.04 0.02 0.01 0

0.05 0.03 0.02 0.01 0 0

0.01 0.01 0 0 0 0

0 0 0 0 0 0

2.94 1.78 1.05 0.61 0.36 0.19

19.57

30.77

26.33

10.8

5.88

3.96

1.81

0.47

0.29

0.01

99.89

Z-numbers of the compound types:: +2: paraffins; 0: 1-ring naphthenes; -2: 2-ring naphthenes; -4: 3-ring naphthenes; -6: alkylbenzenes; -8: Naphthenebenzenes; -10: di-naphthenebenzenes; -12: naphthalenes; -14: naphthocycloparaffins; -18: triaromatics.

Table 1. Z-CN Matrix representation of the feedstock, wt% (NOISE Analysis)

2 0 -2 -4 -6 -8 -10 -12 -14 -18 total

10

8

Wt,%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


6

4

2

0 5

10

15

20

25

Carbon Number

Figure 1. Wt% distribution of the compound types of the feed The main compound types are mono-and di-ring naphthenes and paraffins. The upper limit of the carbon number is 25.

3. Catalysts Ni and Mo were deposited on Ultra Stable Y (USY) zeolites from Zeolyst International. The composition of the catalysts is given in Table 2a

Catalyst

NiO, wt%

MoO3, wt%


SiO2/Al2O3


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

mole ratio 1

4.12

14.81

30

2

3.69

14.73

12

Table 2a. Catalyst composition The annular catalyst basket of the reactor was filled with 4 g of NiMo/USY catalyst (8.8 ml), diluted with 142 g of alpha alumina (82 ml). The catalyst and alpha alumina were crushed to a size between 710 and 850 µm (-20, +25 mesh). The absence of diffusion limitations was checked by means of a number of conversion tests. The texture properties of the USY catalysts shown in Table 2b were determined by the BET method. Catalyst 1

Catalyst 2

494.5

402.3

Diameter, nm

7.2

7.4

Pore volume, ml/g

0.35

0.25

Surface area, m2/g

Table 2b. USY textural properties

The acidity of the USY catalysts (characterized by NH3-TPD) is displayed in Figure 2. The types of acid sites were determined by pyridine-adsorbed Fourier-transformed infrared (Py-FTIR) The experiments were conducted on a Magna560 FT-IR instrument. The results are shown in Table3 .


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12

Catalyst 1 Catalyst 2 10

8

Intensity

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


6

4

2

0 100

150

200

250

300

350

400

450

500

550

o

T, C

Figure 2. Catalyst NH3-TPD spectrum

Total acidity (200 oC)

Strong acidity (350 oC)

B, mmol/g

L, mmol/g

B, mmol/g

L, mmol/g

Catalyst 1

0.16

0.10

0.11

0.05

Catalyst 2

0.22

0.21

0.13

0.12

Table 3. Catalyst Bronsted and Lewis Acid distribution Both the total acidity and the strong acidity of catalyst 2 are higher than those of catalyst 1.

4. Experimental Program 4.1. Experimental Setup The experiments were performed in a 1.0 L Robinson-Mahoney reactor with stationary annular catalyst basket (Figure 3) and complete mixing of the gas and liquid phases. To control vortexes the catalyst basket has internal and external baffles. The rotating shaft is equipped with two impellers which direct the fluid into the center of the annulus at the top and bottom and through the fixed catalyst bed. The liquid exits through a line that has to reach up to the ceiling of the reactor, so as to avoid a separate gas zone with different mixing characteristics. The temperature was measured by thermocouples and monitored by temperature controllers. The pressure was controlled by a back pressure regulator. The effluent of the reactor, consisting of gas and liquid,



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

was separated at the reactor pressure by means of a cyclone coupled with a demister. The gas phase was analyzed on-line by means of a gas chromatograph (GC) with a TCD and a FID detector. The main gas flow was cooled for condensing out heavy fractions and then scrubbed by means of a 20 wt% sodium hydroxide solution to remove hydrogen sulfide before venting. The liquid product was cooled and flashed under ambient conditions. The light gases, dissolved in the liquid phase, were desorbed and collected in a glass burette. The TCD/FID GC was used to analyze and quantify the light gases. The liquid product was collected, measured and analyzed off-line by means of a GC-MS.

Figure 3. Flow scheme of the experimental setup

4.2. Catalyst sulfidation and hydrocracking operating conditions The catalysts were pre-sulfided by a mixture of H2S/H2 containing 10% H2S. The temperature of the reactor was increased from 120 C to 330 C at 30 C/h and then kept at 330 C for 5 hours for


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complete sulfidation. The dependency of the conversion on temperature, space time based on the molar VGO feed rate, W/Fo and the H2/VGO feed ratio were investigated for both catalysts. The Tables 4a and 4b show the operating conditions.

Run

1

2

3

4

5

6

7

8

Reactor temperature,℃

370

385

385

385

385

385

385

395

Total Pressure, bar

82

82

82

82

82

82

82

82

51.2

19.2

22.2

27.2

51.2

51.2

96.1

51.2

10

10

10

10

10

13

10

10

W/Fo, Kg cat· h/Kmol VGO H2/VGO Molar Feed ratio

Table 4a. Operating conditions for catalyst 1

Run

1

2

3

4

5

6

7

Reactor temperature,℃

360

370

385

385

385

385

385

Total Pressure, bar

82

82

82

82

82

82

82

51.2

51.2

27.2

51.2

51.2

51.2

22.2

10

10

10

10

13

16

10

W/Fo ,Kg cat· h/Kmol VGO Molar H2/Feed ratio

Table 4b. Operating conditions for catalyst 2

4.3. Analysis of the gaseous effluent The gaseous effluent was analyzed on-line by a gas chromatograph with TCD and FID detector. The TCD was mainly used to determine the concentration of H2 in the gas effluent. The column was a Hayesep D column, the carrier gas contained 8.5% Hydrogen in Helium. The detector temperature was 250C. The FID was used to analyze the composition of the hydrocarbon fraction from C3 to C8 on a RTX-100-DHA (100m;0.25mm;0.5um) 100% dimethyl polysiloxane column and with Helium as carrier gas. The detector temperature was maintained at 290C The hydrocracking operation was monitored by the on-line analysis of the gas effluent using the



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GC with TCD/FID detectors. Light components such as n-propane, i-butane, n-butane, i-pentane and n-pentane were selected to check the evolution of the reactor conditions towards the steady state.

4.4. Analysis of the liquid product The liquid product was collected and analyzed off-line using a GC-MS with temperature programming. More than 175 components were identified and quantified. They formed the basis for the application of the Robinson method that leads to the characterization of the liquid product in terms of 10 types of compounds per carbon number, as shown in Table 51, 2. In addition the method was calibrated by comparing the feedstock analysis with that provided by the NOISE analysis.

5. Influence of the process variables on conversion and product distribution 5.1. Composition of the hydrocracked product The Z-matrix versus carbon number distribution for the experiments was obtained by combining the composition of the gaseous effluent with that of the liquid product. By way of example the wt% distribution of the total hydrocracking product is shown in Table 5 for run 7 of Catalyst 1.

SATURATES -2

MAR

0

C3

2.220

0.000

C4

1.321

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

1.321

C5

2.267

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

0.000

2.267

C6

2.414

1.571

0.000

0.000

0.002

0.000

0.000

0.000

0.000

0.000

3.987

C7

1.818

3.611

0.000

0.000

0.325

0.000

0.000

0.000

0.000

0.000

5.754

C8

2.858

5.439

0.234

0.000

0.770

0.000

0.000

0.000

0.000

0.000

9.302

C9

2.435

5.848

0.889

0.000

0.851

0.000

0.000

0.000

0.000

0.000

10.023

C10

1.830

4.468

1.682

0.023

0.654

0.879

0.000

0.000

0.000

0.000

9.537

C11

1.751

3.364

2.777

0.155

0.462

0.988

0.000

0.000

0.000

0.000

9.497

C12

1.796

2.027

2.976

0.377

0.416

0.785

0.000

0.029

0.000

0.000

8.406

C13

1.426

1.962

3.766

0.780

0.508

0.520

0.082

0.047

0.008

0.000

9.100

C14

1.225

1.529

2.865

1.068

0.381

0.313

0.154

0.032

0.021

0.008

7.597

C15

1.098

1.319

2.257

1.240

0.348

0.219

0.237

0.061

0.045

0.000

6.824

C16

0.865

1.020

1.468

1.140

0.308

0.155

0.186

0.065

0.077

0.000

5.285

C17

0.575

0.721

0.948

0.782

0.239

0.098

0.111

0.058

0.062

0.000

3.593

C18

0.554

0.577

0.769

0.623

0.226

0.049

0.067

0.047

0.043

0.000

2.954

C19

0.304

0.286

0.380

0.287

0.085

0.019

0.023

0.033

0.018

0.000

1.434

C20

0.134

0.120

0.168

0.121

0.044

0.007

0.009

0.005

0.004

0.000

0.611

0.000

-6 0.000

-8 0.000

-10 0.000


-12

TAR

+2

0.000

-4

DAR

CN/Z

0.000

-14 0.000

-18

Total

0.000

2.220

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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


C21

0.039

0.044

0.059

0.046

0.020

0.002

0.003

0.002

0.002

0.000

0.217

C22

0.005

0.006

0.008

0.006

0.002

0.000

0.000

0.000

0.000

0.000

0.028

C23

0.005

0.011

0.012

0.011

0.004

0.000

0.000

0.000

0.000

0.000

0.044

Total

25.906

34.958

21.257

6.661

5.646

4.034

0.871

0.379

0.279

0.008

100.000

Table 5-Detailed compound or class type distribution(wt%) of hydrocracked products of run 7 for catalyst 1 using the Robinson method

5.2. Balances of the Experimental Runs The total mass balances of each run was better than 4%, generally of the order of 2.5%.The mass and molar balance of carbon and hydrogen were of the same order of magnitude.

5.3. Conversions The conversion of the VGO is defined in terms of the change in the mean value of the carbon number distribution of the components and of the standard deviation. The mean provides a good measure of how much the distribution has shifted towards lighter components as a consequence of

the

cracking,

whereas

the

standard

deviation

provides

an

indication

of

any

expansion/contraction in the distribution. The feed (C5-C25) has a carbon number distribution with a mean value of 13.49 and standard deviation 3.88, meaning that about 68% of the C- distribution is within the 13.49 ± 3.88 range. The 50% percentile is at carbon number 13.49. By way of example in run 6 with Catalyst 1 the mean of the carbon distribution of the product is 11.68 ± 3.66. The feed 50% percentile mark (13.49) shifted to the 68.8% percentile, meaning a conversion of 37.6% [(0.688-0.5)/0.5] with a relative negative expansion of the distribution of -5.76% [(3.66-3.88)/3.88].

6. Influence of the operating conditions 6.1. Influence of the H2/VGO ratio on the product distribution



18

H2/Feed=10(x=81%) H2/Feed=13(x=91.6%)

16

H2/Feed=16(x=92.4%) Feed

14 12

Wt,%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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10 8 6 4 2 0 0

5

10

15

20

25

Carbon Number

Figure 4. Carbon number distribution of feed and product for varying H2/VGO ratio. Catalyst 2. T=385C and W/F°=51.2

Figure 4 illustrates the carbon number distribution at 385 C for 3 different conversions obtained with catalyst 2. The results are very close, indicating that the liquid is saturated with H2 and that dehydrogenation on the metal sites is not the rate determining step, a property also of industrial catalysts and operation and yielding a more favorable product spectrum.

6.2. Influence of temperature on conversion


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90

Catalyst 1 Catalyst 2

80

70

Conversion,%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


60

50

40

30

20 350

355

360

365

370

375

380

385

390

395

400

o

T, C

Figure 5. VGO conversion as a function of the operating temperature for Catalysts 1 and 2 at space time W/F°=51.2 and H2/VGO=10 The conversion at three different temperatures for Catalyst 1 and Catalyst 2 are shown in Figure 5. The activity of Catalyst 2 is significantly higher than that of Catalyst 1. The carbon number distributions of the total hydrocracked products as well as the distributions of the different compound types are quite similar with those in the above-mentioned paragraph and are not shown here.

6.3. Influence of the W/F° ratio on the conversion and product composition Figure 6 shows the VGO conversion as a function of space time, W/F°, for both catalysts and confirms that catalyst 2 is much more active than catalyst 1.These experiments were sequentially designed to provide maximum information on the product spectrum and the kinetics of hydrocracking.



90

catalyst 1 catalyst 2

80 70

Conversion,%

60 50 40 30 20 10

20

40

60

80

100

W/F

Figure 6. Conversion versus W/F° ratio at T=385 C and H2/VGO=10 for the two catalysts.

o

W/F =19.2(x=12.2%) o W/F =22.2(x=20.6%) o W/F =27.2(x=22.2%) o W/F =51.2(x=35.4%) o W/F =96.1(x=44.4%) Feed

10

8

6

Wt,%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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4

2

0 0

5

10

15

20

25

Carbon number

Figure 7. Catalyst 1-Evolution with conversion of the C-number distribution of the hydrocracking product. T=385 C and H2/VGO=10.


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18 o

W/F =22.2(x=41.0%) o W/F =27.2(x=54.6%) o W/F =51.2(x=81.0%) Feed

16 14 12 10

Wt,%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


8 6 4 2 0 0

5

10

15

20

25

Carbon number

Figure 8. Catalyst 2 - Evolution with conversion of the carbon number distribution of the hydrocracking product. T=385 C and H2/VGO=10.

Figures 7 and 8 show the carbon number distribution for runs with increasing W/F° and conversion for both catalysts. The peak of the carbon number distribution shifts to a lower range as the VGO-conversion increases, while the distribution narrows. At the higher conversions the production of C3, C4 and C5- components rapidly rises. Figs 9, 10, 11 and 12 show the evolution of the compound type distribution with conversion. They provide quantitative information on the reaction scheme. The relative abundance of the compound types with Z=2, 0 and -6 (paraffins, mono-ring naphthenes, alkylbenzenes) increases with space time, W/F° and conversion, while that of the compound types with Z=-2,-4,-8 and -10 (2-ring naphthenes; 3-ring naphthenes; Naphthenebenzenes and di-naphthenebenzenes) decreases as the W/F° and the conversion increase. At the highest conversion, obtained with catalyst 2, the 3-ring naphthenes have almost completely disappeared and the 2-ring naphthenes have strongly decreased. As a consequence, the amount of 1-ring naphthenes and of the paraffins has significantly increased with respect to the values in the feed.



3.5 o


3.0

2.5

Wt,%

2.0

1.5

1.0

0.5

0.0 0

5

10

15

20

25

Carbon Number

Figure 9. Catalyst1. Evolution with conversion of the distribution of paraffins (Z=2) in hydrocracked VGO. T=385 C and H2/VGO=10

6.0 5.5 o

W/F =22.2(x=41.0%) o W/F =27.2(x=54.6%) o W/F =51.2(x=81.0%) Feed

5.0 4.5 4.0 3.5

Wt,%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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3.0 2.5 2.0 1.5 1.0 0.5 0.0 0

5

10

15

20

25

Carbon Number

Figure 10. Catalyst 2. Evolution with conversion of the carbon number distribution of paraffins (Z= +2) in hydrocracked VGO.T=385 C and H2/VGO=10 The total amount of paraffins, produced from the ring scission and dealkylation of the naphthenes, increases with conversion, but at the highest conversion the C5-C7 paraffins


Page 15 of 28

themselves crack into C3-C4 paraffins, a less desirable feature.

1.0 o


0.8

Wt,%

0.6

0.4

0.2

0.0 0

5

10

15

20

25

Carbon Number

Figure 11. Catalyst 1. Evolution with conversion of the distribution of alkylbenzenes (Z= -6) in hydrocracked VGO. T=385 C and H2/VGO=10 Figure 11 reflects how the alkylbenzenes with higher C-number gradually decrease with conversion and how the peak in the C14-C17 range disappears.

5.0 4.5

W/Fo=22.2(x=41.0%) W/Fo=27.2(x=54.6%) W/Fo=51.2(x=81.0%) Feed

4.0 3.5 3.0

Wt,%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


2.5 2.0 1.5 1.0 0.5 0.0 0

5

10

15

20

Carbon Number


25


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 16 of 28

Figure 12. Catalyst 2-Evolution with conversion of the carbon number distribution of 2-ring naphthenes (Z=-2)in the hydrocracked products.T=385 C and H2/VGO=10

Figure 12 shows how the 2-ring naphthenes decrease with increasing conversion.The maximum of their distribution also shifts towards lower C-number.

6.4. Effect of the catalyst acidity on the product spectrum To investigate the effect of the catalyst acidity on the product distribution, two runs with almost the same conversion for catalyst 1 (Run7) and catalyst 2 (Run7), conducted at the same temperature, H2 partial pressure and H2/VGO ratio, are selected (viz Table 6.a and 6.b). Catalyst

o Run W/F ,Kg.h/Kmol

T,oC Pressure,bar

Catalyst 1

7

96.1

385

82

10

44.4

-3.35

Catalyst 2

7

22.2

385

82

10

41.0

-0.77

H2/VGO Conversion,%

Expansion,%

Table 6.a. Comparison of reaction conditions and results for Run 7 with catalyst 1 and Run 7 with catalyst 2

Catalyst

Run No. Mean Carbon Standard deviation Conversion,%

Catalyst 1

7

11.24

3.75

44.4

Catalyst 2

7

11.4

3.85

41.0

Table 6.b. Conversion and expansion of the carbon distribution


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The carbon number distribution of the hydrocracking products of catalyst 1 and catalyst 2 is shown in Figure 13. The distribution of the compounds with carbon number above 13 is very similar for both catalysts but their amount is significantly lower than in the feed. The amount of compounds with carbon number 6 and 7 obtained with catalyst 2 is higher than that of the catalyst 1, reflecting the higher cracking activity of catalyst 2. Yet, the C3 yield with Catalyst 1 is higher than that of catalyst 2, perhaps because the conversion is slightly higher.

Catalyst 1(x=44.4%) Catalyst 2(x=41.0%) Feed

10

8

6

Wt,%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


4

2

0 0

5

10

15

20

25

Carbon Number

Figure 13. Carbon number distributions of the hydrocracked VGO for catalyst 1 and catalyst 2. H2 pressure=82 bar, T=385 C and H2/VGO=10. Conversions respectively 44% and 41% Figure 14 compares the wt% distribution of the various compound types of the feed and hydrocracking products of catalyst 1 and catalyst 2 for close global conversions of respectively 44 and 41%. For the two catalysts the wt% of compound types with Z=2 (paraffins) and Z=0 (mono-ring naphthenes) has increased through the operation, while that of compound types with Z= -2 and -4 (2-and 3-ring naphthenes) has decreased. There is little difference between the two catalysts for the amount of higher naphthenes in the product, but for catalyst 1 the abundance of



mono-ring naphthenes is higher than that observed with catalyst 2, while that of the paraffins, the ultimate products of the sequence, is lower. The higher acidity of catalyst 2 has apparently more weight in the lower carbon range. 40,00 35,00

relative abundance,%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 18 of 28

30,00 25,00 20,00

catalyst 1

15,00

catalyst 2

10,00

feed

5,00 0,00 2

0

-2

-4

-6

-8

-10 -12 -14 -18

Z value

Figure 14. Comparison of compound type distributions obtained with catalyst 1 and 2. H2 pressure=82 bar, T=385 oC and H2/VGO=10. Conversions of respectively 44 and 41%

7. Kinetic modeling The information collected in this study permits the development of a detailed kinetic model for VGO hydrocracking using the single event concept, introduced by Froment and co-workers3-7. The elementary steps involved in the hydrocracking are: Acid catalyzed steps: - Carbenium ion (de)protonation - Alkene protonation - Isomerization by hydride and by methyl shift In hydrocracking these steps are very fast and reach equilibrium. The following steps are kinetically controlled: - Isomerization by protonated cyclopropane (PCP) branching - β-scission in acyclic ion - β-scission in exocyclic chain


Page 19 of 28

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


- Endocyclic β-scission - Dealkylation of aromatics - Cyclization These six types of steps are encountered with many, if not all members of the concerned homolog series of components and lead to a huge number of rate coefficients. Reactions on the metal sites, following chemisorption: - Cycloalkane and aromatic (de)hydrogenation - Alkane and alkene (de)hydrogenation. With well designed hydrocracking catalysts these reactions reach equilibrium.

The number of elementary steps occurring in the hydrocracking of the hydrogenated VGO dealt with here ranges in a few hundred thousand. Therefore, the reaction scheme leading to the final product spectrum is generated by computer using Boolean relation matrices and characterization vectors3,4. The number of rate coefficients is evidently extremely large, but the fundamental transition state theory for kinetics offers perspectives on how to deal with this problem. In terms of transition state theory the rate coefficient for a given type of elementary step A⇁B is written as:

(1) Within a given type of hydrocracking elementary step the value of the rate coefficient k was shown experimentally to evolve with the structure of the species. This evolution can be formulated by selecting a characteristic of the structure of the reactant and activated complex: the symmetry number. The ratio of these symmetry numbers is called number of single events4,

 σ glR ne =  ‡ σ  gl

  . It is calculated using quantum chemical methods or approximations thereof. 

Factoring out this ratio from the exponential entropy group in (1) leads to:

 σ gRl k = ‡  σ gl 

  k BT    h

 ∆Sˆ o‡   −∆H o‡    exp   exp  RT  R     


(2)


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

The group

Page 20 of 28

is the “single event frequency factor”. It does not depend

upon the structure of the reactant and activated complex and is unique for a given type of elementary step. Its value is calculated from the experimental data.

The change of enthalpy associated with an elementary step is expressed in terms of its heat of reaction8, to be calculated by group contribution methods9-10 and of two additional parameters, A and α, the intrinsic activation barrier and the transfer coefficient. These take on unique values for all the elementary steps of a given type and are derived from experimental data. The model contains 21 independent parameters, to be determined from the experimental data. Their values were estimated by minimizing the Root Mean Square Error, defined as the weighted difference between model and experimental responses. The optimization strategy alternated a series of algorithms including Differential Evolution, Coordinate Descent and Trust-Region methods11-12. The Root Mean Square Error evolved from an initial value of 1.0 x 10-3 to a constant value of 7.1 x 10-5 after 54 trials. Fig. 15 shows a parity plot comparing the mass yields of the various subclasses.


Page 21 of 28

Figure 15. Model vs. Experimental mass yield parity plot

8. Application to process simulation The kinetic model is a powerful tool for the simulation of commercial hydrocracking units. The following figures, all related to the simulation of a commercial hydrocracking operation comprising a 3-zone adiabatic fixed-bed reactor with inter-zone quenching , were generated using the current implementation of the kinetic model available in the ProMax® Process Simulator13. They were obtained using a commercial VGO feedstock (VGO 2-Table 7) heavier than the one chosen for the present experimental and kinetic modeling study. The latter is less representative for the feeds used in commercial hydrocracking on a high activity catalyst like cat 2. Its application would lead to the conversion of valuable fractions and to excessive yields of gas. By virtue of the single event approach the kinetic parameter values obtained in the present study are valid also for the hydrocracking of this commercial feed, in fact for any VGO type feedstock,

1 2 3 4 5 6 7 8 9 10 11 12 13

0.000 0.000 0.000 0.000 0.000 0.000 0.001 0.009 0.040 0.084

0.000 0.000 0.000 0.000 0.000 0.001 0.009 0.040 0.093

0.000 0.000 0.062 0.093 0.093 0.084

0.000 0.000 0.000 0.000 0.000 0.000

0.000 0.000 0.000 0.018 0.062 0.036 0.077 0.144

0.275 0.329 0.207 0.291

0.036 0.036 0.077 0.329


Total

Dicycloalkane-Monoaromatics

Monocycloalkane-Diaromatics

Monocycloalkane-Monoaromatics

Triaromatics

Diaromatics

Monoaromatics

Tricycloalkanes

Dicycloalkanes

Monocycloalkanes

Isoalkanes

provided the same catalyst is used.

Carbon Number

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.02 0.44 0.51 0.53 1.03


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 Total

0.144 0.199 0.207 0.467 0.473 0.371 0.325 0.325 0.344 0.881 0.617 0.862 0.387 0.334 0.241 0.411 0.393 0.360 0.291 0.262 0.292 0.000 0.000 0.000 0.000 0.000 0.000 8.320

0.077 0.144 0.199 0.207 0.467 0.473 0.371 0.325 0.325 0.344 0.253 0.617 0.365 0.387 0.334 0.241 0.411 0.393 0.360 0.291 0.262 0.000 0.000 0.000 0.000 0.000 0.000 6.989

0.144 0.329 0.379 0.179 0.291 0.319 0.321 0.490 0.519 0.654 0.253 0.551 0.458 0.651 0.748 1.068 0.411 0.393 0.499 0.465 0.466 0.000 0.000 0.000 0.000 0.000 0.000 9.920

0.379 0.207 0.179 0.269 0.319 0.605 0.402 0.490 0.519 0.325 2.221 0.881 0.458 0.365 0.387 0.334 0.241 0.552 0.499 0.465 0.466 0.000 0.000 0.000 0.000 0.000 0.000 10.563

0.199 0.379 0.269 0.319 0.321 0.490 0.519 0.654 0.253 0.551 0.862 0.651 0.334 0.241 0.411 0.393 0.360 0.291 0.262 0.292 0.349 0.000 0.000 0.000 0.000 0.000 0.000 8.737

0.321 0.371 0.490 0.667 0.325 0.344 0.253 0.617 0.862 0.387 0.334 0.241 0.411 0.393 0.360 0.291 0.262 0.292 0.741 0.557 0.708 0.708 0.000 0.000 0.000 0.000 0.000 11.037

0.325 0.654 0.253 0.551 0.365 0.365 0.387 0.748 1.068 0.638 0.552 0.499 0.465 0.466 0.546 0.349 0.741 0.557 0.708 1.812 1.067 0.000 0.000 0.000 0.000 0.000 0.000 13.116

0.179 0.269 0.319 0.473 0.371 0.490 0.667 0.654 0.253 0.551 0.458 0.651 0.748 1.068 0.638 0.393 0.360 0.291 0.466 0.546 0.349 0.000 0.000 0.000 0.000 0.000 0.000 10.672

0.605 0.402 0.325 0.325 0.253 0.253 0.551 0.458 0.365 0.651 0.748 0.241 0.411 0.393 0.360 0.291 0.262 0.292 0.349 0.557 0.557 0.000 0.000 0.000 0.000 0.000 0.000 8.649

Page 22 of 28

0.179 0.269 0.319 0.321 0.325 0.667 0.325 0.344 0.253 0.881 0.458 0.365 0.387 0.334 0.638 0.552 0.499 0.465 0.291 0.546 0.349 0.000 0.000 0.000 0.000 0.000 0.000 8.767

2.55 3.22 2.94 3.78 3.51 4.38 4.12 5.11 4.76 5.86 6.76 5.56 4.42 4.63 4.66 4.32 3.94 3.89 4.47 5.79 4.87 0.71 0.00 0.00 0.00 0.00 0.00 96.77

Table 7. Commercial VGO Feedstock (VGO 2) Characterization

Figure 16 represents a typical temperature profile for such a commercial 3-zone adiabatic reactor. The catalyst distribution is designed to limit the temperature rise in the first bed to 35 K, to avoid less favorable product distributions.


Page 23 of 28

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 16. Temperature profiles in a 3-zone adiabatic fixed-bed reactor

Figure 17 shows the evolution of the yields of various commercial fractions with conversion. These are obtained from the summation of the model-calculated yields of individual components or compound types within relevant carbon numbers. The trends are completely in line with commercial experience. The simulation illustrates the very high activity of the catalyst that causes valuable fractions, even of this heavier feed, to be partly converted into lighter and less valuable products. Figure 18 shows reactor profiles of the yield of fractions consisting of saturated species. The cyclohexanes are rapidly cracked in the first two beds, generating paraffins, which in turn crack further into lower molecular weight paraffins. Within a certain type of elementary step each component of a homolog series is formed from higher members but decomposes into a lower member, as in any consecutive reaction scheme (with contributions here from parallel steps also). That is why the figure distinguishes, by way of example, between the yields of lumps of paraffins with carbon numbers above and below C17. The same is done for the mono- and di-aromatics in Figure 19, showing how the rapid hydrogenation of aromatics in the first two zones forms



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

saturated rings subject to cracking

Figure 17. Evolution of the yields of commercial fractions with conversion in a 3-zone adiabatic fixed-bed reactor


Page 24 of 28

Page 25 of 28

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 18. Evolution of the yields of saturated species with conversion in a 3-zone adiabatic fixed-bed reactor



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 26 of 28

Figure 19. Evolution of the yields of aromatic species with conversion in a 3-zone adiabatic fixed-bed reactor

9. Conclusions The present paper deals with the generation of a comprehensive and detailed data base for the effect of the operating conditions and catalyst acidity on the conversion and yield patterns of the hydrocracking of a heavy oil fraction. It required adequate experimental facilities and considerable analytical efforts but allowed for the derivation of a detailed fundamental Single Events kinetic model for VGO hydrocracking and its further implementation in the ProMax® Process simulator. Such a model is obviously an important tool for the scale-up of bench and pilot data supporting the front-end engineering design of grass-root hydrocracking units as well as for improving the performance of existing ones. It can also contribute to providing quantitative insight in the development of new catalysts.


Page 27 of 28

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Acknowledgement This work was supported by Bryan Research & Engineering. Dr Tao Zhang is grateful to the Chinese Scholarship Council, the National Natural Science Foundation of China and the China National Petroleum Corporation (U1362203).

Symbols A = single-event frequency factor E = activation energy, KJ/mol Fo = liquid flow rate of feedstock, Kmol/h h = Planck’s constant, kJ·h k = rate coefficient kb = Bolzmann’s constant, kJ/(molecule·°C) R = universal gas constant, kJ/(mol·K) T = temperature, K

∆H°‡ = enthalpy of activation, kJ/mol ∆S°‡ = intrinsic entropy of activation, kJ/(mol·K) σRgl = global symmetry number of reactants σ‡gl = global symmetry number of activated complex

References (1) Robinson Jr, C. J.; Cook, G. L., Low-resolution mass spectrometric determination of aromatic fractions from petroleum. Anal. Chem (Washington,D.C., USA). 1969, 41, 1548-1554. (2) Robinson, C., Low-resolution mass spectrometric determination of aromatics and saturates in petroleum fractions. Anal. Chem (Washington, D.C.,USA). 1971, 43, 1425-1434. (3) Baltanas, M. A.; Froment, G. F., Computer generation of reaction networks and calculation of product distributions in the hydroisomerization and hydrocracking of paraffins on Pt-containing bifunctional catalysts. Comp. Chem.Eng. 1985, 9, 71-81. (4) Froment, G. F., Single event kinetic modeling of complex catalytic processes. Catalysis Reviews 2005, 47, 83-124. (5) Froment, G., Kinetic modeling of hydrocarbon processing and the effect of catalyst deactivation by coke formation. Catalysis Reviews 2008, 50, 1-18. (6) Kumar, H.; Froment, G. F., Mechanistic kinetic modeling of the hydrocracking of complex feedstocks, such as vacuum gas oils. Ind. Eng. Chem. Res. 2007, 46, 5881-5897. (7) Kumar, H.; Froment, G. F., A generalized mechanistic kinetic model for the hydroisomerization and hydrocracking of long-chain paraffins. Ind. Eng.Chem. Res. 2007, 46, 4075-4090.



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(8) Evans M. and Polanyi.M., Inertia and driving force of chemical reactions. Trans. Faraday Soc. 1938, 31, 0011-0023. (9) Marrero, J. and Gani, R., Group-contribution based estimation of pure component properties. Fluid Phase Equilibria 2001, 183-184, 183-208. (10) Hunter, E.P.L., Lias, S.G., Evaluated Gas Phase Basicities and Proton Affinities of Molecules. J Phys. Chem. Ref. Data. 1998, 27, 3. (11) Rao, S.S., Engineering Optimization, 4th Ed., Wiley, New Jersey 2009. (12) Biegler, L.L., Nonlinear Programming, MOS-SIAM Series on Optimization, SIAM, Philadelphia 2010. (13) ProMax® Process Simulator Manual, BRE Group Ltd., Bryan 2014.


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