Catalysts for upgrading solvent refined lignite
by Nam Kyun Kim
A thesis submitted in partial fulfillment of the requirements for the degree of Doctor of Philosophy in
Chemical Engineering
Montana State University
© Copyright by Nam Kyun Kim (1982)
Abstract:
The solvent refined lignite (SRL), made at the University of North Dakota Process Development Unit,
was a solid having a nominal melting point of 160°C. The SRL was pulverized and mixed with a donor
solvent, tetralin. The SRL to tetralin ratio of 1:1 was selected to pretreat in a high pressure and
temperature reactor. The optimized reactor conditions were a reaction temperature of 475°C, an initial
hydrogen pressure of 2000 psig and a retention time of 40 minutes. Under these conditions
approximately 97% of the SRL was dissolved in tetralin. The resulting solution was used to test the 27
developmental catalysts.
The catalysts were developed by impregnating on the γ-alumina the 3 active metals; MoO3, CoO, and
WO3, each at 3 levels. The effect of these factors on upgrading of the SRL was evaluated in terms of
denitrogenation, desulfurization, and hydrocracking. The multiple linear regression analysis showed
that the metal compositions for the best overall catalytic performance were 9.5% MoO3, 4.3% CoO,
and 4% WO3 (% of carrier weight).
A model was developed based on the results of scanning electron micrographs to explain some of the
physical characteristics of the catalysts. The disadvantage of the incipient wetness method used in
metal impregnation was explained, and the preferable pore structure and distribution were suggested. CATALYSTS FOR UPGRADING SOLVENT REFINED LIGNITE
by
Nam Kyun Kim
A thesis submitted in partial fulfillment
of the requirements for the degree
of
Doctor of Philosophy
in
Chemical Engineering
MONTANA STATE UNIVERSITY
Bozeman, Montana
October 1982
3 )3 %
K Sbias
d o jo • <3ii
APPROVAL
of a thesis submitted by
Nam Kyun Kim
This thesis has been read by each member of the thesis committee and has been found
to be satisfactory regarding content, English usage, format, citations, bibliographic style,
and consistency, and is ready for submission to the College of Graduate Studies.
Date
Chairperson, Graduate Corhmittee
Approved for the Major Department
L c ,
Date
/ /
Head, Major Department
Approved for the College of Graduate Studies <
Date
Graduate Dean
iii
STATEMENT OF PERMISSION TO USE
In presenting this thesis in partial fulfillment of the requirements for a doctoral degree
at Montana State University, I agree that the Library shall make it available to borrowers
under rules of the Library. I further agree that copying of this thesis is allowable only for
scholarly purposes, consistent with “fair use” as prescribed in the U.S. Copyright Law.
Requests for extensive copying or reproduction of this thesis should be referred to Uni­
versity Microfilms International, 300 North Zeeb Road, Ann Arbor, Michigan 48106, to
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tation in and from microfilm and the right to reproduce and distribute any abstract in any
format.”
Signature
Date
iv
ACKNOWLEDGMENT
The author would like to thank the Department of Energy and Associated Western
Universities, Inc. for the financial support that made this research possible. Special thanks
are given to Dr. Lloyd Berg, director of this research, and Dr. F. P. McCandless for thenguidance. The moral support and encouragement of the staff of the Chemical Engineering
Department are gratefully acknowledged.
Appreciation is extended to Lyman Fellows for his fabrication of the research equip­
ment, Ms. Ahce Brekke for proofreading, Andy Bnxt for the preparation of SEM, Dr. E.
Abbott for the NMR, and Mrs. M. C. Wagner of Katalco for the analyses of surface areas
and pore distribution. A special thanks goes to Dr. R. Lund, who helped with the statistical
design.
FinaUy, I wish to express my gratitude to my wife, Sook, son, Daniel (9), and daughter,
Nancy (7). Without their encouragement and patience, I could never have been a graduate
student.
V
TABLE OF CONTENTS
Page
APPROVAL PAGE................................................................................................................... ...
STATEMENT OF PERMISSION TO USE..................................................
ui
ACKNOWLEDGMENT........... ....................................................................
iv
TABLE OF CONTENTS...........................................
v
LIST OF TABLES...................................................................
LIST OF FIGURES........... ..................................
ABSTRACT........................................................................................................
yii
viii
xi
INTRODUCTION..................................................
I
BACKGROUND.................................................................................................................
3
Lignite...........................................................
Solvent Refined Lignite.........................................................................................
Chemical Structure of Lignite..........................................................................
Liquefied Lignite...........................................................................................................
Catalytic Upgrading.....................................
Catalyst.............................................
Trickle Bed R ea c to r....................................................................................................
Research Objective......................................................................................................
5
5
10
12
15
17
20
20
EXPERIMENTAL.............................................
. Feedstock.....................................................................................................................
Preparation of C atalysts............................
Continuous Trickle Bed R eactor................................................................................
Operation of Continuous Trickle Bed R ea c to r.........................................................
Analytical Procedure........... .........................................................................................
RESULTS AND DISCUSSION............................................................................
Preparation of Feed S o lu tio n .............................................
Performance T ests.........................................................................................................
Effect of Metal Compositions on Upgrading.............................................................
Development of A Model for Catalyst........................................................................
22
22
24
30
32
33
34
34
46
50
75
vi
Page
SUMMARY AND CONCLUSIONS....................................................
92
RECOMMENDATION FOR FUTURE STUDY .................................
94
LITERATURE C ITE D ..................
APPENDICES........................................................................
A. Surface Area and Pore Distribution of Catalyst Carriers
B. Sample Calculation of Pore Volumes............................
C. ASTM D-86 Distillation Data.........................................
D. Modeling of Katalco Carrier............... ..
100
101
119
123
127
VU
LIST OF TABLES
Table
Page
I. Composition and Characteristics of Feed and Products................. ..............
9
II. Analyses of SRL and North Dakota Lignites..................................................
23
III. Analysis o f A sh............................................................................ ..
...........
24
IV. Systematic Preparation of 27 Catalysts; .........................................................
26
V. Summary of the First 15 Batch R u n s .............................................................
35
VI. Effect of Water Addition on Catalytic Perform ance.....................................
37
VII.
SRL Dissolubility at Various Operating Conditions......................................
40
VIII.
SRL Dissolubility at Extended Operating Conditions..................................
42
IX. Catalytic Performances: KT-14 vs. Blank C arrier.........................................
47
X. Nitrogen Removal With KT Series C atalysts...............................................
49
XI.
Summary of Catalytic Performance............ ................
XII. Analysis of Variance for Three Factors at Three Levels.. . . ' ..............
52
60
XIII. Multiple Regression Analysis for Denitrogenation....................................
61
XIV. Analysis of Variance for Desulfurization as Response Variables.................
70
XV. Multiple Regression Analysis for Desulfurization.........................
71
XVI.
Multiple Regression Analysis for Gasoline Y ield..........................................
76
XVII.
Multiple Regression Analysis for Heavy Oil Y ield.........................................
77
Comparison of Various Catalysts.................................
83
XVIII.
XIX. Pore Volume and Surface Area Following 10% MoO3,
4% CoO, and 8% WO3 Impregnation on KatalcO Carrier...............................
85
XX. Pore Volume and Surface Area Following 10% MoO3,
4% CoO, and 8% WO3 Impregnation on Nalco A Carrier...............................
86
viii
LIST OF FIGURES
Figure
Page
1. Coal fields in the Northern Great Plains province...............................................
4
2. Schematic flow diagram for 50 Ib/hr PDU with mass rates and
run conditions............................................. .................................... .................
g
3. Chemical precursors of coal.....................................................................................
11
4. Representative partial structures of c o al...............................................................
13
5. Comparison of lignite with other sources of hydrocarbon
..........................
14
6 . A theoretical molecule and thermal breakup of coal (Wiser)...............................
16
7. Approximate reaction routines for the HDS, HDN, and HDO from
five membered heterorings in the presence of catalyst.........................................
18
8 . A 3-D representation of 3 3 experimental design...................................................
27
9. Electric furnace and auxiliary equipmentfor sulfidation of catalyst...................
29
10. Trickle bed reactor arrangement............................................................................
31
11. Effect of water addition in wt% denitrogenation and wt%
desulfurization............................................... ........................................................
38
12. Effect of reactor temperature and retention time at two levels of
hydrogen pressure on SRL dissolubility in te tra lin ............... ..............................
41
13. Effect of reactor temperature on SRL dissolubility in tetralin
at two levels of retention time.................................................................................
44
14. Effect of hydrogen pressure on SRL dissolubility at two different
temperature levels....................................................................................................
45
15. Effect of retention time on SRL dissolubility in tetralin at 475°C
and 2000 psig of H2 ......................................................
45
16. Catalytic performance o f KT-14 against blank base in denitro­
genation ..................................................................................
48
17. ASTM D-86 distillation, feed and 8-hr composite product of blank
carrier...............
51
ix
Figure
Page
18. Catalytic performance in denitrogenation as a function o f MoO1
ennr.Anrra+irm
3
54
19‘ C^alytic performance in denitrogenation as a function of CoO
WO3, and MoO3 concentrations...................................
’
20. Catalytic performance in denitrogenation as a function of CoO
concentration................................................
2 1. Catalytic performance in denitrogenation as a function of MoO3
WO3, and CoO concentrations...................................
’
22. Catalytic performance in denitrogenation as a function of WO3
concentration..............................................
23. Catalytic performance in denitrogenation as a function o f MoO3
CoO5and WO3 concentrations...................................
24. Catalytic performance in desulfurization as a function of MoO3
concentration..............................................
25. Catalytic performance in desulfurization as a function of CoO
WO3, and MoO3 concentrations........................ ...................
26. Catalytic performance in desulfurization as a function of CoO
concentration.......................................................
27- Catalytic performance in desulfurization as a function of MoO3
WO3, and CoO concentrations i ........................
28. Catalytic performance in desulfurization as a function of WO3
concentration................................................
29. Catalytic performances in desulfurization as. a function of MoO3
CoO5and WO3 concentrations...................................
30. Catalytic performance in hydrocracking as.a function of CoO1
WO3 , and MoO3 concentrations........... ............................
3 1. Catalytic performance in hydrocracking as a function of MoO3
WO3, and CoO concentrations.......................................
’
32. Catalytic performance in hydrocracking as a function of MoO3
CoO, and WO3 concentrations.................
’
55
56
57
58
59
63
64
65
66
67
68
72 .
73
74
X
Figure
33. Scanning electron photomicrographs of various catalysts:
(A) Katalco blank carrier..............................................
(B) KT-14 with 10% MoO3,4% CoO, and 8% WO3 on KataIco
c arrie r........................................................................
(C) KT-14 after 8-hr run .........................................................
(D) Union Carbide Linde 13X with 10% MoO3 ,4% CoO' and
8% WO3 ..........................................................................
(E) Ketjen LA-3P with 10% MoO3 ,4% CoO 5and 8% WO3
(F) Nalco A with 10% MoO3 ,4% CoO, and 8% WO3. . .
(G) Harshaw CoMo 0401 with 9% MoO3 and 3% CoO...........
34. Three dimensional arrangement of basic granules for Katalco
carrier.......................................
Page
78
78
78
79
79
80
80
81
35. Effect of catalyst pore diameter on denitrogenation
36. Surface area distribution vs. pore diam eter.............
37. Distribution OfMoO3 vs. pore diameter
89
38. Representative structures of pore
90
xi
ABSTRACT
The solvent refined lignite (SRL)5 made at the University of North Dakota Process
Development Unit, was a solid having a nominal melting point of 160°C. The SRL was pul­
verized and mixed with a donor solvent, tetralin. The SRL to tetralin ratio of I I was
selected to pretreat in a high pressure and temperature reactor. The optimized reactor con­
ditions were a reaction temperature of 475°C, an initial hydrogen pressure of 2000 psig
and a retention time of 40 minutes. Under these conditions approximately 97% of the SRL
was dissolved in tetralin. The resulting solution was used to test the 27 developmental
catalysts.
catalysts were developed by impregnating on the 7 -alumina the 3 active metals'
MoO3, CoO, and WO3, each at 3 levels. The effect of these factors on upgrading of the
SRL was evaluated in terms of denitrogenation, desulfurization, and hydrocracking. The
multiple linear regression analysis showed that the metal compositions for the best overall
catalytic performance were 9.5% MoO3, 4.3% CoO, and 4% WO3 (% of carrier weight).
A model was developed based on the results of scanning electron micrographs to
explain some of the physical characteristics of the catalysts. The disadvantage of the
incipient wetness method used in metal impregnation was explained, and the preferable
pore structure and distribution were suggested.
I
INTRODUCTION
Currently the United States consumes close to 80 quadrillion Btu (quads) per year,
importing increasing amounts of oil to meet its needs. The United States by 1978 already
imported almost 8.5 million barrels of oil per day-over 17 quadrillion Btu per year. It is
generally recognized that this level of imports is unhealthy for the U.S. economy. Every
worker in the United States consumes 1.55 billion Btu per year for all purposes of employ­
ment [ I ] . If employment is to continue to grow, energy will have to be supplied to the
society in increasing amounts from coal. Production of oil and natural gas is declining at an
annual rate of about 4 to 5%. Nearly 45% of the petroleum that is currently produced has
been obtained by water flooding and other secondary recovery techniques applied to
mature fields [ 2 ].
In the present United States environment of decreasing availability of petroleum and
natural gas, coal is a natural candidate for the raw material for liquids and gases. The
reasons for this are that the United States has more energy available in the form of coal
than in the combined sources of petroleum, natural gas, oilshale, and tar sands. The use of
coal for energy will certainly increase in the United States during the next several decades.
Utilization of domestically abundant coal, both for the production of power and as a feed­
stock in synthetic fuels production, will require processing operations on a large commer­
cial scale [3 ].
The transportation sector demands exclusively liquid fuels, the residential and com­
mercial sectors depend heavily on gaseous fuels, and three quarters of the energy used by
industry are constituted as liquid and gaseous fuels [ 2 ] . Consequently, the conversion of
2
coal to gaseous and liquid fuels in commercial quantities is vital to ensuring the availability
of fuel in conventional forms for the major users.
Coal gasification and liquefaction processes were pioneered during the 1920s and
1930s in Germany. They constitute the basis for much of today’s technology. The accom­
plishment of Friederich Bergius [4] brought him the Nobel Prize for chemistry in 1931.'
His direct hydrogenation of coal at elevated temperature (430° C or 806° F) and pressures
(3,000-10,000 psig) led to the production of gasoline and aviation fuel. At the same time
Mathias Pier and co-workers found sulfur resistant, coal-hydrogenation catalysts that
reduced the severity of the environment for liquefaction while improving conversion effi­
ciency. The production of synthetic fuels from coal will have justification due to the con­
venience of using liquids and gases and due to the ways in which transportation and domes­
tic systems have been developed [5].
BACKGROUND
The United States Geological Survey estimated that the lignite shares 6% of the
demonstrated coal reserve base (437 billion tons). The lignite in the Northern Great Plains
(NGP) occurs in relatively thick seams ranging from 5 ft to more than 100 ft and typically
with overburden from 50 to 200 ft. The ratio o f “overburden volume to lignite volume” is
most favorable when compared with that o f higher ranking coal [ 6 ] ..
Most significant is the fact that the NGP surface mineable reserve base totals 82.3 bil­
lion tons or 58% of the national surface mineable reserve base. More than 95% of these
resources are lignite and subbituminous (classified as low rank coals) occurring in the Fort
Union and Powder River Regions and the Bull Mountain Field (Figure I).
Nearly 100% o f NGP coal mined since the 1960s has been used to generate electricity,
but the future potential for the production of synthetic fuels and chemicals is increasing.
Since 1971, plans for construction of seventeen separate mine-mouth coal synfuel plants in
the NGP have been announced. These plans include eight separate gasification plants, nine
liquefaction plants, and two in-situ gasification pilot operations [ 7 ].
\
The first commercial scale synthetic fuels project, the Great Plains Gasification Plant
(GPGP) of Oliver County, North Dakota, has already begun on July 25, 1980 and is
scheduled for full gas production by the end of 1984 [ 8 ]. Project Lignite, equivalent to
the SRC-I, was awarded to the University of North Dakota in 1972 for the purpose of
determining the appropriate technological approach to the conversion of lignite to clean
fuel. The original plan was expected to extend the two stage conversion o f lignite to liquid
fuel (equivalent to the SRC-II) with the solvent refined lignite (SRL) as an intermediate
solid fuel. The first stage to convert lignite to the SRL has been successfully accomplished,
F ort
,U nion
k I R e g io n
M ontana
B u llS>v—
M o u n ta in
F ie ld
S o u th
D a k o ta
'owde
I v e r J <5
R e g io
m ile s
N o r th
D a k o ta
IIIlIB l i g n i t e
Wyoming
Figure I . Coal fields in the Northern Great Plains province.
fc==3 s u b b it u m in o u s c o a l
5
but the second stage that catalytically hydrotreats the SRL to the distillate fuel was not
implemented [9].
The objective of this research is to catalytically upgrade the SRL to liquid fuel for
immediate industrial use or to clean distillate suitable for conventional refinery feedstock.
Lignite
Lignite is a brownish-black coal that is intermediate in coalification between peat and
subbituminous coal. According to the classification system adopted by the American
Society for Testing and Materials (ASTM), the lignite is the lowest rank of coals in terms of
calorific value (less than 8300 Btu per pound on a moisture, mineral-matter-free basis) and
carbonaceous content (47 to 59%) [10].
While the use of lignite to generate electricity will predominate other uses, a strong
potential also exists for the conversion of lignite to synthetic fuels. The Great Plains Gasifi­
cation Plant is designed to convert the North Dakota lignite into pipeline quality synthetic
natural gas (SNG) having about 977 Btu per standard cubic foot (scf). Approximately
137 million Scf per day of SNG and other byproducts such as anhydrous ammonia, tar,
oil, phenols, and naphtha will be produced by processing 22,000 tons per day of lignite.
Lignite displays unique properties; (I) its high reactivity evidenced by spontaneous com­
bustion, ( 2) non-coking and non-swelling nature with high permeability upon heating, and
(3) excellent sulfur absorbent qualities [ 6 ].
Solvent Refined Lignite
.
The Project Lignite Process Development Unit [11] has a normal design capacity of
50 pounds of lignite feed per hour and produces light liquids and gases in addition to
approximately 15 pounds per hour of SRL having a melting point of 150 to 205°C (300 to
400° F).
6
Lignite as received is pulverized to slurry with solvent. The slurry is pressurized, pre­
heated and reacted at a selected temperature (normally at 434°C or 814°F) and pressure
(2500 psig) in a reducing gas environment. The products are then separated as gases, liquids,
and SRL from the unreacted coal and mineral matters. The flow rate of solvent to the
slurry mixing tank is controlled by the signals from an orifice in the solvent feed line. The
flow out of the tank is controlled by the slurry level in the tank.
Two dissolvers (or reactors), R-IA and R-IB, are made from 18-ft lengths of 4 7/8inches OD by 3 7/8-inches ID Incoloy 800 tubing. The inlet of each reactor is at the bot­
tom and the outlet at the top, with another outlet at the center. Thus any multiple of
10 ft lengths up to 40 ft can be assembled. Consequently, residence times for the slurry
can be varied four-fold at a constant feed rate.
The gas-slurry mixture then goes to a series of separators at high, intermediate, and
low pressures. The vapor products consist primarily of unreacted carbon monoxide and
hydrogen, carbon dioxide, hydrogen sulfide, ammonia and light hydrocarbons.
The liquid products consist mainly of water, process solvent, and lighter products.
Light end column F-2 is operated at 10 to 15 psig. The overhead product is light oil, essen­
tially a stabilized naphtha* consisting of hydrocarbons ranging from about C5 to perhaps
as high as 350°F to 450° F boiling point. The bottom product is recycled to the slurry mix
tank.
Mineral separation consists of a system o f vessels and pumps designed for the high
temperature extraction of solvent refined lignite plus solvent from the mineral matter and
unconverted lignite using toluene as a diluent. The toluene-slurry mixture is then fed to the
vacuum flash drum F-I via the settling tower V-8 (18 inch diameter by 12 ft high) in which
the terminal velocity of the settling particles is greater than fluidizing solvent (toluene) and
SRL velocity.
The bottoms from the toluene flash vessel V-9 is fed through preheater E -II and into
the vacuum flash drum F -I. The bottoms from F-I is the SRL product.
A typical PDU operation is shown in a flow schematic for Run M-33C (Figure 2). In
this run, 47 Ib/hr of average 36.7% moisture lignite was processed with 407 scf/hr of gas
containing 50% H3 and 50% CO, 91 Ib/hr of recycle in the liquid-solid separation system.
The rough material balance is shown,as well as pressures and temperatures in the important
vessels. The 100 lb of moisture and ash free (MAF) lignite produced 57.4 lb of SRL and
Ught organic Uquids, consuming 7.8 lb of CO, 0.04 lb of H3, and 2.38 lb of water with a
wt % MAF coal conversion of 93.4 [12].
A typical composition and characteristics o f SRL product from Zap Ugnite are shown
in Table I. GeneraUy it has been customary to classify the quaUty of the coal products in
terms of solubiUty classes. OUs are hexane soluble fractions, whUe asphaltenes are terms
used for benzene or toluene soluble materials. The benzene insoluble material is preasphaltene (some prefer the term asphaltols or polar compounds) which are soluble in pyridine.
The primary, product of coal is pyridine soluble,but benzene insoluble. This fraction is sub­
sequently converted to both benzene-soluble and hexane-soluble species through Uquefaction process.
Hexane-soluble materials (oUs) average about 200-300 in molecular weight. They have
Uttle or no functionaUty. Asphaltenes, on the other hand, are predominantly mono-func­
tional compounds. They consist of phenols and basic nitrogens. Molecular weights range
from 300 to 700. Asphaltols have multiple functionaUty. They consist o f polyphenols (up
to 5 OH/molecule) and multiple basic nitrogens. Molecular weights range from 400 to
2000 or greater [13] . Conversion in this case is defined as conversion o f coal to material
soluble in pyridine. This fraction is not found in petroleum and suggests a considerable
basic difference between coal and petroleum structure. Low hydrogen content and high
16.68
P r o d u s t gas to
therm al o x id iz e r
S-2
W aste g as
T: 75
Feed gas
R eactor
T: 802
P : 2500
I
CO
46.79
91.08
R e c y c Le
T:607
P : 2500
S -3
T: 440
P : 300
2.55
S-4
T: 88
P :60
Bo t t o m s
>
(80% w a t e r )
20.10
, Toluene
accum ulator
0.0
V -8
T:44:
P ; 30C
—> ( 1 2 . 6 0
22.43
L ight
o ils
V -8 b o tto m s so I id l
J & heavy li q u i d s
Cold 'J S x h a u s t g a s e s
L iquids
1 4 . Ov!/ t r a p
Vacuu m b o t t o m s ( m . p . = 3 0 0 ° F )
S o lv e n t accum ulator
3.26
12.32
0.42
0.32
14.67
0.0
T o t a l 2. 41 .2 8
2 4 1 . 28 T o t a l
Run No . : M-33C (.6/2 3 / 7 7 )
U n its : T (=te m p e r a t u r e ) ,° F ;
Mass r a t e s , I b s / h r
P (= p ressu re), psig
Figure 2. Schematic flow diagram for 50 Ib/hr PDU with mass rates and run conditions.
9
Table I
Composition and Characteristics of Feed and Products
F e e d Gas
CO
H2
( P r o j e c t L i g n i t e PDU)
V o l . ■I
W t. %
24.7
81. 9
75.2
17. 8
M a t e r i a l B a l a n c e f o r Gas C o m ponents
Vol% I n
V o l% Out
I b / h r In
75.2
H2
66.2
2.13
CO
24.7
14.4
9.78
CO2
—
13.2
H2 S
—
—
0.2
—
CH4
—
4 .5
—
.
C2H6
1 .1
—
0.3
c 3h 8
NH3
0 .1
U ltim ate A nalysis o f M a te ria ls
L ig n ite
S tartin g
R ecycle
Charged
S olvent
S olvent
C
45.22
89.03
83.60
H
6.43
8.1 1
9.14
N
0.64
0.1 2
0.2 0
S
0.4 5
2.23
1.09
0*
41.45
0.51
5.97
A sh
5.81
0.0
0.0
I b / h r Out
1.87
5.81
9.91
0.18
1.11
0.58
0.33
0.04
Vacuum
Bottom s
80 . 2 0
5 .20
0 .98
0,. 9 0
4, . 4 6
8,. 2 6
Deashed
SRL
87.42
5.67
1.07
0.98
4.86
0.0
* By d i f f e r e n c e
P r o p e r t i e s o f P r o d u c t SRL
G ra d ie n t Bar M elting P o in t , F
P y r i d i n e S o l u b l e s , wt% a s h - f r e e
S p e c ific G ravity
H eat o f C om bustion, B t u / l b
M easured* C a l c u l a t e d
321I
83. 5
100
1 . 2 !8
1.25
1 4 , 330
15,990
* F - I vacuum b o tto m s
L ig n ite:
N o r t h A m e r i c a n C o a l C o . , Z a p , N .D .
S c r e e n s i z e 9 0 % - 200 m e s h , 1 0 0 5 - 6 0 m e s h
M o i s t u r e 31 .5 %
10
heteroatom content, compared to petroleum, make coal somewhat intractable with con­
ventional refining technology.
The solvent-refining process consists mainly of conversion o f insoluble coal fo the
pyridine-soluble, toluene-insoluble fraction of SRL. The net result is an increase in aromaI
ticity and some bond breakage, loss of about 20% of the original carbon as gases and vola­
tile liquids, and possible reduction of oxygen, nitrogen^ and sulfur.
Chemical Structure of Lignite
It is generally agreed that coal may originate primarily from plants. Through a
sequence of evolutionary changes the primary products of the original decomposed plant
materials become transformed. The first product is humic acid. Then, the humic acid is
transformed eventually into peat, lignite, subbituminous coal, bituminous coal, and finally
anthracite [14].
The United States coals consist of primarily vitrinite, usually 80% or more. The com­
position of this vitrinite is believed to be the result of the coalification of either cellulose
or lignin structures, which constitute the majority of the plant components [15]. Some of
the chemical precursors of coal are shown in Figure 3.
Ligmte is considered to be a crosslinked amorphous polymer, with mostly mono­
diaromatic aggregates connected by relatively weak cross-links. Generally there have been
two approaches to deducing chemical structure. One way is to break down the coal mater­
ial into recognizable fragments and then put them back into an original structure. An alter­
native approach is direct characterization of solid coal with the use of sophisticated instru­
ments such as IR spectroscopy, NMR spectroscopy and X-ray diffraction. These modem
techniques are also severely handicapped because coal is not crystalline and insoluble.
These advanced techniques such as X-ray scattering have been used in the past and
conflicting inteipretations as to the predominant structure of coal have been reported.
11
CH2OH
C ellu lo se
CH2 OH
CH2OH
L ignins
HO OH
Waxes
23 47
Model o f a hum ic a c i d
Figure 3. Chemical precursors of coal.
12
New instruments have evolved recently that are capable of direct characterization in its
solid form. The most promising of these tools is a solid state CP-C13 NMR developed by
Pines[16].
Wender indicated that the carbon skeleton of coals can be considered as consisting
of hydroaromatic structure with aromaticity increasing from low-rank to high-rank coals.
Figure 4 shows some frames of reference for various ranks of coal [17].
Liquefied Lignite
Coal has chronic problems in utilization. It is a solid of non-uniform composition
inorganic material and environmentally objectionable elements such as sulfur, nitrogen,
mercury, etc. Conversion of lignite to liquids or gases substantially reduces these disadvan­
tages for lignite. The most important chemical change required for this conversion is the
addition of hydrogen (hydrogenation). The amount of hydrogen addition determines the
quality of the synthetic product. A comparison of some representative fuels illustrates in
Figure 5 the scale of hydrogen/carbon mole ratio [14]. Lignite has a lower hydrogen con­
tent than that characteristic of premium quality transportation fuels like gasoline and diesel
oil. It can be seen that there is a long path necessary in the conversion of lignite to high
quality products.
Liquefaction of coals where the liquid products are the main product has been known
for many years since the first work by Bergius. Subsequent development of the process of
coal liquefaction have led to a variety of process conditions for producing liquids. The
term liquid may need to be defined since some products of coal liquefaction are solids at
room temperatures. The degree of conversion can be measured by the amount of material
soluble in a certain organic solvent [ 2 ].
Wiser [18] showed in 1975 a schematic representation of structural groups and con­
necting bridges in bituminous coal. It may consist of layers of condensed aromatic and
13
L ignite
ZCHnOH
Subb itu m in o u s
COOH
OH
H ig h -v o latile
bitum inous
L ow -volatile
bitum inous
A nth racite
0 -
Figure 4. Representative partial structures of coal.
SRL
S y n th o il
& H -co al
D istilla te
Peat
A sp h a lt &
T arsan d
Pe t r o l e u m
Premium
ItProducts
R esid
C o als
H y d r o g e n / c a r b o n m o le r a t i o
T e tra lin
S R C -I H - c o a l F u e l
o il
A n th ra c ite
coal &
SRL
coke
L ig n ite
OTE
H -c o a l n a p h th a
( s y n c r u d e mode)
S R C -II
A rab l i g h t c r u d e
COED
EDS N a p h t h a
1.0
1.5
H y d r o g e n / c a r b o n m o le r a t i o
Figure 5. Comparison of lignite with other sources of hydrocarbon.
2.0
15
hydroaromatic clusters ranging in sizes from one to several rings per cluster with an average
of three rings per condensed configuration. The significance of these theoretical molecules
is the location of a number of relatively weak bonds indicated by arrows which can
account for the easy thermal breakup of coal into smaller more soluble fragments (Figure
6).
Catalytic Upgrading
In the process of the hydrogenation of lignite to produce synthetic liquid some
removal of heteroatoms (sulfur, nitrogen, and oxygen) is also accomplished. Sulfur and
nitrogen contents of lignite are often greater than 1% and oxygen content is sometimes
over 20%. Such heteroatoms are responsible for some of the coal conversion and upgrad­
ing problems. Upgrading process not only improves the heating value of the fuel but also
makes resulting products more environmentally acceptable.
Hydrogenation of liquefied coal is slower than that of petroleum crudes because of
the abundance of the polynuclear aromatic compounds. Oxygen is removed primarily as
carbon dioxide and water with small amounts of carbon monoxide. About 40-50% of the
oxygen and organic sulfur is relatively easy to remove. It is believed to be the result of
exchange of OH or carbonyl oxygen by sulfur, due to biological activity in the sediment
[19]. The remaining sulfur is much more resistant to removal and is probably present in
heterocyclic ring structures.
Removal of S, N, and O from SRL under reducing conditions and in the presence of
an industrial catalyst is associated with elimination of hydrogen sulfide, ammonia, and
water. Prior to these reactions, C-Y (Y = S, N, or 0 ) bonds may have to be broken, and the
fission of one of these bonds may be the rate controlling step [ 20 ].
Heterocyclic S, N, and O containing rings are well known for their high resistivity to
removal. Ring saturation may be required for N containing compound, while there is some
16
,-4-c-I
N -C -H
C a ta ly tic h y d ro c ra c k in g
& h y d ro g e n a tio n
i—
IW -H
H
o
H
N -C -H
N -C -N
Figure 6. A theoretical molecule and thermal breakup of coal (Wiser).
17
experimental evidence for the HDS with or without preliminary heteroring hydrogenation
[ 21 ] .
The basic routes for hydrodesulfurization (HDS)5 hydrodenitrogenation (HDN)5 and
hydrodeoxygenation (HDO) of heterocyclic compounds are shown in Figure 7 as suggested
by Furimsky [22]. The hydrogenation of the heteroring is an equilibrium process affected
by the concentration of hydrogen [23].
Hydrocracking is extensively practiced commercially in petroleum refining to produce
high quality gasoline, jet fuel, diesel, high quality lubricant [24,25,26,27]. Some of the
commercially proven catalytic hydrocracking methods are the Standard Oil of Indiana
Ultracracking Process and Union Oil Unicracking Process. These processes can tolerate
feedstocks with a nitrogen content of as high as 0 .3% [28,29].
Under trickle bed catalytic hydrotreating conditions the denitrogenation is always
accompanied by other reactions such as hydrogenation, hydrocracking, desulfurization,
depxygenation, coking, and demetallization.
Catalyst
It has long been known that the rates of chemical reactions can be accelerated by
small amounts of alien material. Such material is termed a catalyst and it is defined as a
substance which increases the rate at which a chemical reaction approaches equilibrium,
without being consumed in the process [30]. An appropriate catalyst plays a key role in
removing sulfur, nitrogen, and oxygen simultaneously as gaseous hydrogen sulfide, am­
monia, and water from the syncrude oil.
The carrier, quite often alumina, refers to a major catalyst constituent that serves as
a base or binder for the active metals and promoters. A carrier may be catalytically active
or inert. The major function of a earner is to provide a large surface area so that catalyti­
cally active metals can be spread out or dispersed as a monolayer, if possible [31,32,33].
18
CH=CH2
CH=CH2
+
Y=S:
H2Y
H2S
B en z o th io p h e n e
Y=N:
NH 3
N
In d o le
Y=O:
H2 O
B en zo fu ran
Figure 7. Approximate reaction routines for the HDS, HDN, and HDO from five membered
heterorings in the presence of catalyst.
19
The transition metal oxides and their mixtures with elements of group IV B and V B
of the periodic table are of great interest as selective oxidization catalysts [ 34 ].
Molybdena and tungstate promoted by cobalt or nickel are well known active agents
for their characteristic activities of HDS, HDN, and HDO [35,36]. Activity and selectivity
of different catalyst systems were investigated by Qader [37]. Cobalt sulfide on silica
(low) alumina was found to be the most active catalyst in the hydrocracking of polycyclic
aromatic hydrocarbons. Pelleted tungsten sulfide was the most successful early hydro­
cracking catalyst [38].
The hydrocracking catalyst has dual functions. They are (I) cracking of high molecu­
lar weight hydrocarbons, and ( 2) hydrogenation of the unsaturates formed either during
the cracking step or already present in the feedstock. A balance of hydrogenation and
hydrogenolysis activity is vital to maintain catalytic activity [36A]. This balance can be
controlled by the appropriate combination of metals and methods of catalyst preparation.
Beuthef et al. reported that in most commercial applications the molybdenum-cobalt in
atomic ratios varies from 0 . 1: 1.0 to 1.0 : 1.0 but the best activity was observed for ratios
around 0.3:1.0 [39].
Popov [40] and Bliznakov et al. [41] demonstrated that a mixture of WO3 and
MoO3 had a higher activity for oxidation of methanol to formaldehyde than either of the
pure oxides.
Effect of silicon dioxide concentration on physicochemical properties of hydrocrack­
ing catalyst were investigated by Perezhigina et al. [42]. Addition of SiO2 to a CoMoO4 /
Al2O3 catalyst increased its cracking and isomerization ability, conversion and. rate of
iso-to-normal hydrocarbons in the final products. Further, mechanical strength of the
catalyst was increased by a factor of 1.5, while a slight decrease of hydrodesulfurization
activity was experienced.
20
Trickle Bed Reactor
The trickle bed reactor is a device in which a liquid phase and a gas phase flow cocurrently downward through a packed bed of catalyst material to promote desired reactions.
When the fixed bed of catalyst is used with a liquid or mixed liquid plus vapor reactant
under operating conditions, the trickle bed reactor is an appropriate choice. Here the
liquid reactant is fed at the top o f the reactor to distribute evenly over the area of the cata­
lyst bed. Each catalyst particle is wet with the liquid feed as it trickles down through the
bed. The reacting gas penetrates the liquid film and reacts on the surface of the catalyst
[43].
Use of trickle bed reactor in the petroleum industry involves the processing with
hydrogen of various petroleum materials. The HDS, HDN5HDO, and catalytic hydrocrack­
ing ofheavy or residual oil stock to upgrade the quality has been economically tested using
this type of reactor. Ross [44] pointed out that effective distribution of liquid over the
catalyst in trickle-phase hydrogenation was the key factor affecting the overall reactor
efficiency. The velocity dependence of liquid distribution over the catalyst is a compli­
cated function of the initial distribution, bed-packing arrangement, catalyst particle geom­
etry, wettability of the liquid on the catalyst, local, gas velocity, etc. Since the catalyst
particles retain a finite amount of liquid both on the external surface and in the pores,
the variables that influence the liquid distribution determine its residence time distribu­
tion. Therefore, the bench scale trickle phase reactor is suited for the initial product
evaluation, rather than the derivation of the reaction kinetics [45 ].
Research Objective
This research aims to statistically evaluate the effect of the metal concentrations (Mo,
W, and Co) on the upgrading of solvent refined lignite (SRL) using Katalco alumina-silica
carrier. Upgrading of SRL to liquid fuel for immediate industrial use or to clean distillate
suitable for conventional refinery feedstock includes reduction of nitrogen and sulfur contents and increase in clean products recovered in the ASTM D-86 distillation.
The research plan constitutes three parts. Phase I consists of a pretreatment of SRL to
convert it to a manageable liquid feed at room temperature. It includes an optimization of
operating conditions for the batch reactor to determine in which conditions SRL promises
to make the most dissolution in tetralin. Phase II consists of statistical design of catalysts
having three metals, each at three levels. Performance of each catalyst was evaluated in a
continuous trickle bed reactor. Phase III consists of the development of a geometric model
for the catalyst to explain the physical characteristics and the performance of the catalyst.
22
EXPERIMENTAL
The SRL was dissolved in a solvent, 1,2,3,4-tetrahydronaphthalene (tetralin), using a
batch autoclave with hydrogen at elevated temperatures and pressures. This pretreatment
was necessary for preparation of liquid feed, since the SRL as received was a solid. The
optimum operating conditions were established for the maximum dissolution of SRL in
tetralin and the common feedstock was prepared using these conditions. Twenty-seven
catalysts were fabricated using three-factorial design approach. The design is a completely
randomized design and the levels of the factor considered are fixed levels. The catalysts
were tested in continuous trickle bed reactor to evaluate the performances. The method
of chemical analyses is described later in this section.
Feedstock
The SRL was received from the University of North Dakota [45]. A representative
analysis of SRL (PDU Run M-33) and the lignite mined near Gascoyne, N.D., is listed in
Table II. The SRL is a hard-brittle solid with an incipient melting point of about 90° C at a
barometric pressure of 26.5 in. Hg. The SRL was made with 50:50 hydrogen-carbon mon­
oxide syngas at a reactor pressure of 2500 psig and a maximum dissolver temperature of
820° F (438°C).
The X-ray fluorescent analysis of ash is shown in Table III [45]. Ash is the noncom­
bustible mineral matter when lignite is burned under specified conditions of temperature,
time, and atmosphere (ASTM D-3174). Any of these constituents may deteriorate the cata­
lytic activity.
Table II
Analyses of SRL and North Dakota Lignites
Q u an titativ e
A nalysis
Zap
L ig n ite
S olvent
SRL
Coal
A sh
W ater
T o tal
64.70
6.56
28.74
100.00
P y rid in e S olubles
Elem.
A nal.
Carbon
H ydrogen .
N itro g en
S u lfu r
O xygen(by d i f f .)
Ash
T otal
A s p h a l­ wt% a sh
tene
Wt % u n c o n v e r t e d c o a l
Test
wt% p r e a s p h a l t e n e s
Wt% a s p h a l t e n e s
wt% m a lte n e s & d i s t . o i l
T otal
SRL f r o m
F -I h ot.
(M-39)
6.4
92.3
0.28
1.02
100.00
Gascoyne
L ig n ite
54.84
8.48
36.68
100.00
99.72
46.29
5.83
0.48
0.29
40.55
6.56
100.00
87.74
6.15
0.87
0.67
3.55
1.02
100.00
1.02
0.28
27.88
38.34
32.48
100.00
SRL f r o m
F -I h ot.
(M- 33)
2.5
97.32
0.18
100.00
100.0
38.28
6.83
0.51
0.71
45.19
8.48
100.00
85.73
5.98
0.80
0.94
6.37
0.18
100.00
0.18
21.20
36.46
42.16
100.00
24
Table III
Analysis of Ash
(L ignites)
.
Origin ( N o r t h D a k o ta )
Loss on ignition (808° C)
SiO2
Al2 O3
Fe2 O3
TiO2
P3 O3
CaO
MgO
Na2O
K2O
SO2
Total
Wt %
Zap
Gascoyne
0.4
20.2
10.5
10.0
0.5
0.6
26.7
6.8
6.7
0.4
17.2
100.0
37.6
56
5.6
0.8
06
20.5
6,9
4.1
0.3
10.8
100.0
Preparation of Catalysts
A complete block of catalysts has been fabricated using three active metal compo­
nents impregnated at three levels on a commercial carrier, Katalco extrudates (serial no.
81-6731). Water soluble salts of these metals were selected for the incipient wetness
method. Three metal salts chosen are listed below.
1. Ammonium molybdate: (NH4)6Mo7O24 -4H20 (F.W. = 1235.9) with 81.4% essay
InMoO3 (M.W. = 143.94)
2. Cobalt nitrate: Co(M )3 )2 -6H2 O (F.W. = 291.050) with 99.5% purity. The M.W.
o f CoO is 94.7326.
3. Ammonium Metatungstate: S(NH4)2O- 12W03 -7H2 O (F.W. = 3168.67) with 99%
purity. The M.W. of WO3 is 231.8482.
The Katalco carrier prior to metal impregnation has the following physical properties
[46].
Surface Area: 223 sq m/gm
Pore Volume: 0.933 cu cm/gm
25
Pore.Diameter: 169A
The pore volume distribution data measured by the relative amounts of nitrogen absorbed
or desorbed at different absolute pressures is attached in Appendix A.
The sequence of metal loading on a blank carrier has been consistent in the order of
Mo5 Co, and W throughout the preparation of 27 catalysts. The following procedure served
as a general guide for catalyst preparation [47].
1. Drying and calcining the carrier
2. Contacting the carrier with impregnating solution
3. Removing the excess solution
4. Drying and calcination
5. Activation
The carrier was dried at IlO 0C in an oven and calcined at 500°C in a muffle for
8 hours to obtain the net carrier weight. The volume of the metal solution was determined
for each metal impregnation so as to minimize the excess solution of the next metal salt.
Generally, the amount of excess solution varied from 30% to 50%. A systematic prepara­
tion of 27 catalysts is summarized in Table IV and a 3-D representation is shown in Figure
8. A specific example of calculation for the pore volume is presented in Appendix B.
The final metal oxide concentration of a given catalyst by a single impregnation may ■
be calculated from the pore volume and the concentration of solution neglecting selective
adsorption [47]. For example, an approximate percent MoO3 is obtained on the total
weight basis as follows, if the pore volume is 1.1 cu cm/gm and MoO3 concentration in
solution is 10.1%.
1.1 X 0.101
i + L i x 0.1101 x
100 = 10%
The wet catalyst after draining off the excess solution was dried by placing under the
draft hood for approximately two hours. Following the oven-drying at IlO0C for about
26
Table IV
Systematic Preparation o f 27 Catalysts
2%WO3 1— D&C{— 4 2
-|2%CoO|—
8
%W0 3 H D & C H
4 2
2
8
H14%WO 3 1—fD&Cl—I 3 2 14
2%WO3 1—[P&C]—|. 4 4
H 4 %Mo 0 3 ^—[d [--|4%Co O]—fp|--j '8 %WO3 1—[D&c]—| 4 4
- f lliw o s H D & c H
2
8|
4 4 141
rj 2%W03HD&CH 4 6
-| 6 %CoOKDV-l 6%W03]—|D&C|—I 4 6
Ml4%W03HD&clH 4
6
2|
8
14
2 %wo 3 1—Td ^ c T—11 o 2 ~~2
H2%Co O H d'H
8%W 03H d &c H T o 2
8
Ml4%WO3MD&cHT0 2 14|
B lank
— D - - 1 0 %Mo Q3 —[D}-|4%CoO|—[D
C arrier
H 2%W03HPD&CHTq' 4
2|
8%WO3HD&CT-|10 4
8
4l4%W03[—[D&U}—fTo~ 4 14
H 2%WO3HD&CHT0 6
- 6 %CoO—
8
%W0
3
l—D&Cj—110
6
21
81
H14%WO3H d & C H 1 0 6 1 4 1
H 2 %W0 3 t- P g & C H l 6 2 2|
2%Co q M d I--I 8%wo3HD&c]- r i 6 2
8
-| 14%WO3 1— D&C — 16 2 14
H 2%W 03H d &c H 1 6
4
2
-j 1 6 %MoO^—(D}f4%CoO|- 4 D ^ -T 8%W0^f—fD&C}—116 4
8
H14%WO3 1—fD&C]—116 4 14
"6%Co 01M d M
2%W03|- fD & c H l6 6
D r D r y i n g a t I l O 0 C f o r 12 h r s
l ..^ w o 3 H P i £ K ^ Cr C a l c i n a t i o n a t 5 0 0 ° C f o r 8 h r H 1 4 %W0 3 M D & c M 16
N ote:
%M e t a l O x i d e i s
d efin ed :
6
6
2l
81
14|
(Wt. MOx ZWt c a r r i e r ) *100
27
« 1 ------- I -
- 4 - - I __
------------ r
I
Dark b a l l s : R e p l i c a t e s
MoO 3
Figure 8. A 3-D representation of 33experimental design.
28
12 hours the catalyst was calcined at 500 C for 8 hours. The calcined catalyst was cooled
to room temperature in a desiccator and weighed for final metal concentration.
A batch o f catalyst (60 ml) was loaded in a I inch diameter sulfiding tube with a con­
centric thermowell for temperature measurement. The tube was heated to 450° C by an
electric furnace while a stream of 10% H2S in hydrogen passed through the catalyst for
12 hours at 2 to 3 ce per second rate (Figure 9).
The molybdenum, cobalt, and tungsten catalysts were activated by reduction and
sulfidation in order to obtain an active catalyst for most reactions in which they are
employed (except for oxidation-type reactions). The reduction proceeds gradually and
smoothly from the Mo+* (W+6) state to lower valance states (MoS2 or WS2) in a stream
of 10% H2S in hydrogen at 450°C [35]. Cobalt oxide is believed to reduce to Co9S8
form. Even when presulfiding is not employed, the catalysts become sulfided during pro­
cessing due to the hydrogen sulfide liberated from the reaction. However, incomplete
sulfiding produces catalyst with inferior performance and life.
Sulfided catalysts are known to have better activity than the catalyst in oxidic form.
Schuit et al. [48] explained that the bond strength of S-H is significantly lower than that
of O-H and a distance between the surface and molybdenum ion is increased by replace­
ment O with S. This may decrease the interaction of it electrons from the N heterorings
with molybdenum, resulting in less favorable conditions for surface poisoning by nitro­
gen bases.
The spent gas from the sulfiding tube was scrubbed by a series of impingers with a
20% NaOH solution. The low concentration (20 to 150 ppm) OfH2 S gas can be identified
by a distinctive odor similar to rotten eggs.
Exposures of 800 to 1,000 ppm may be fatal in 30 minutes. H2 S does not combine
with the hemoglobin of the blood; i t ’ s asphyxiant due to paralysis of the respiratory
center [49,50].
y
Powe r s t a t
20%NaOH
so lu tio n
Figure 9. Electric furnace and auxiliary equipment for sulfidation of catalyst.
30
Continuous Trickle Bed Reactor
The bench scale trickle bed reactor used for this research was fabricated by the
Chemical Engineering Department of Montana State. University (MSU). A schematic dia­
gram of the trickle bed reactor and its auxiliary equipment is shown in Figure 10.
The reactor consisted of a I -in. I D. (35 in. long) schedule-80 Inconel pipe. The top of
the reactor was welded with a % in. stainless steel cross which allowed the fitting of a
33-in. piece of stainless steel tubing, which served as a thermowell, and the fitting of two
feed ports, one for the feed solution and another for hydrogen. The reactor was fitted into
the bore of a 6-in. ti.D. 3-ft long aluminum block wrapped with three sets of ceramic bead
encased NiChrome heating wires. The power to each heating wire was individually con­
trolled by a Powerstat variable autotransformer. A chromel-alumel (type K) thermocouple
wire was placed in the %-in. stainless steel thermowell at the center of the catalyst section.
The temperatures were read by Cole-Palmer digital thermometers (model 8520-40).
To load the catalyst the empty reactor was secured upside down by a vise. The upper
space of the reactor was loaded with 175 ml of 1A-In. Norton Denstone, inert ceramic pel­
lets [51], and followed by 25 ml of 1/8 in. Denstone mainly to preheat and uniformly
distribute the feed solution. Sixty ml of catalyst diluted with 60 ml of 1/8 in. Denstone
were loaded [52]. Additional 45 ml of 1/8 in. Denstone filled the remaining space of the
reactor. Finally a stainless steel screen, acting as a support, was placed at the bottom of the
reactor before a reactor plug was threaded into the pipe.
The preheated feed solution was metered into the top of the reactor by the use of a
Milton Royal simplex piston pump (Model A MR-1-23) through a 1/8 in. stainless steel
tube. The pumping rate was manually controlled by an attached micrometer. AU the feed
lines and the reservoir were wrapped with flexible heating cords [53] so that the feed
temperature was kept around IOO0C by the use of Powerstats. Technical grade, hydrogen
31
D ig ital
therm om eter
F lex ib le
Therm ow ell
Pressure
relief
valve
Check
valve
NiChrome
R eservoi
Ammeter.
eacto r
B u rette
P ow erstat
l~
NiChrome
heating
w ire
M ilton-R oy
p isto n
Back
pump
pressure
reg u lato r
Vent
A lu m in u m
h eatin g
block
Figure 10. Trickle bed reactor arrangement.
Bypass
valve
‘I n s u l a t i n g
m aterial
Condenser P r e s s u r e
reg u lato r
C old w a t e r
G as-liq u id
sep arato r
20%NaOH
so lu tio n
M icro-
Product
o u tlet
Hydrogen
conta in e r
32
supplied in a cylinder was metered through a pressure regulator by a Brooks thermal mass
flowmeter [54].
Hydrogen and feed solution were passed cocurrently through the pressurized reactor
(1,000 psig) and then to a gas-liquid separator through a Grove back-pressure regulator
equipped with a corrosion-resistive Teflon diaphragm. The exit gases were scrubbed by a
20% NaOH solution and vented. The liquid products were collected in either continuous
or periodic mode from the separator.
Operation of Continuous Trickle Bed Reactor
The reactor loaded with the catalyst to be tested was placed in the aluminum heating
block. The feed solution and hydrogen lines were connected at the top of the reactor and
the separator system was secured at the bottom of the reactor. The thermocouple wire was
inserted into the thermowell. All pipe fittings were sealed using either Teflon tape or anti­
seize compound (Permatex). The entire system was pressurized to a normal operating range
of 1000 psig and checked for leaks using Snoop soap solution. If no leaks were detected,
the system was depressurized. Three Powerstats were turned on to heat the reactor to
425° C. When the reactor reached desired reaction temperature, the entire feed system
(reservoir, burette, feedlines, check valves, and piston housing) was preheated with
NiChrome wires. The hot feed solution was recirculated for a few minutes through the
feed line back into the reservoir. Then the feed line was connected to the reactor for oper­
ation. The reactor was pressurized with hydrogen through the bypass valve attached to the
Brooks flowmeter. When the system approached desired pressure, the bypass valve was
closed and the micrometer was adjusted to keep the hydrogen flow to the reactor at a
10,000 scf/bbl rate. The feed solution valve located at the top of the reactor was then
opened. The pump was started and the time was logged. The feed rates were measured by
timing the liquid level drop in the burette. The feed rate was adjusted by the use of pump
33
micrometer to a liquid hourly space velocity (LHSV) of 1.0 and frequently checked to
ensure a steady flow. The products were collected hourly following the startup, unless
otherwise specified.
After the last sample was obtained, the pump was shut off and the liquid feed valve at
the top of the reactor was closed. The excessive feed solution was drained from the reser­
voir and the feed system was cleaned with solvent. The hydrogen valve was shut off and
the line was disconnected. The Powerstats were shut off and the separator system was
removed after depressurization. The reactor was removed from the aluminum heating
block and the entire content of the reactor was dumped in an orderly manner on the col­
lector plate for further inspection. The reactor was cleaned by brushing with acetone.
Analytical Procedure
Hourly liquid products from all runs were analyzed for nitrogen, and a composite
sample was analyzed for sulfur to represent the run. The ASTM D-86 distillation was
carried out for each composite product. The nitrogen contents were analyzed by the MSU
Chemistry Station-Analytical Laboratory using the macro Kjeldahl method [55,56]. Sul­
fur concentration was determined by the dual unitized quartz tube combustion apparatus
[57,58]. Weight percent denitrogenation (% DN) and weight percent desulfurization
(% DS) were calculated as follows.
Fns - Pns
—
— X 100 = % DN (or % DS)
where Fns = Wt% nitrogen or wt% sulfur of the feed
Pns = Wt% nitrogen or wt% sulfur of the product.
The extent of hydrocracking was determined by ASTM D-86 distillation. Fifty ml of
composite sample representing each catalyst were used as a standard volume. This method
measures the cumulative amount of the distilled product which boils below 700°F (370° C)
or when decomposition begins [59].
34
RESULTS AND DISCUSSION
A total o f eighty-two batch runs was carried out with a Parr Instrument Company
Series 4,000 Pressure Reaction Apparatus for determination of optimum dissolution of
SRL in tetralin and for preparation of common feed solution. A complete block of 27
catalysts was tested using a continuous trickle bed reactor. In addition, nine tests of cata­
lysts were duplicated. Finally five special catalysts with different pore diameters were pre­
pared and tested using a large molecular solute as a tracer. The objective of this part of
the investigation was to see how meaningful is the average pore diameter of a catalyst,
especially when the solute diameter approaches pore diameter.
Preparation of Feed Solution
A mortar and pestle were used to grind chunks of SRL for all runs. The slurry was
prepared by mixing powder SRL in tetralin while heating and agitating with a combination
magnetic stirrer-heater. Two hundred ml of the resulting slurry (at 130°C) were charged to
a 500 ml Parr autoclave. The autoclave cap was then screwed on and the cap bolts were
tightened with a torque wrench. The autoclave was pressurized to 1,000 psig with hydro­
gen and then depressurized to the barometric pressure. This barometric pressure of hydro­
gen was designated as 0 psig of H2 . The autoclave was again pressurized, if necessary, to a
I
desired final pressure with hydrogen. A summary of the first fourteen autoclave runs, is
I
shown in Table V.
The first two runs were made to see if the high temperature Could help dissolving the
powder SRL in tetralin. The SRL to tetralin ratio of 1:2 was used and the air was not
evacuated for these runs. Although the second run (K-2) was extended twice as long in
retention time as the first run under the similar operating conditions, both products were
Table V
Run
no.
Gas
K -I
K -2
-K-3 K-4
K-5
K- 6
K -7
K- 8
K-9
A ir
A ir
H2
H2
H2
H2
H2
H2
H2
K-IO
K -Il
K - 12
K -13
K - 14
Pres­
sure
p sig
Summary of the First 15 Batch Runs
( D i s s o l u t i o n o f SRL)
Temper­
R eten­
SRL
atu re
tio n
T etralin
°C
•time
ratio
hr
1500
1-500
1500
150 0
150 0
150 0
5 30
5 40
475
4 75
475
4 75
475
475
475
2
0 .5
0 .5
0 .5
I
I
I
I
I
I
H2
150 0
4 75
2
I
H2
H2
H2
H2
1500
1500
1500
1500
4 75
475
4 75
475
I
I
I
I
I
I
I
I
0
0
1000
I
2
2
2
2
2
2
3
Remark
Not d is s o lv e d
Not d isso lv ed .
D issolved
D isso lv ed . P roducts
o f K -4 t h r o u g h K - 7
were used f o r
Run K C - 1.
D issolved
. 2 0 ml o f w a t e r w a s
added. D issolved.
1 0 ml o f w a t e r w a s
added. D isso lv ed .
5 ml o f w a t e r was
added. D issolved
• P roducts of K -Il
th ro u g h K -14 w ere
u s e d f o r K C -2
36
messes of slurries, showing no sign of dissolution. The third run (K-3) was carried out in
the same fashion with the pressurized hydrogen instead of air for two hours at 475° C.
The product was a liquid with some sediments at room temperature. This progress allowed
reducing the SRL to tetralin ratio from 1:2 to 1:1. The following four consecutive runs
(K-4 through K-7) were made to prepare feed solution for a continuous trickle bed reactor
run.
The catalyst, KN-A, was tested in Run KC-I using the feed solution decanted from
the composite reservoir. In five hours on stream the product became dark and viscous. The
operating conditions and the nitrogen and sulfur contents of the hourly products are pre­
sented in Table VI. Performance of the catalyst in hydrodenitrogenation is illustrated in
Figure 11. The first sign of catalytic deactivation appeared in the fourth hour on stream,
and from then on it progressively worsened. The yield was 67.5 wt %.
In ah effort to extend the life of the catalyst, a small amount of water was added to
the slurry as practiced in the conversion of ethylbenzene to styrene. The Run K-8 was
made with a 20 ml of water added to the 200 ml of the slurry under the same reactor con.
e
/
ditions as Runs K-4 through K-7. The resulting product showed an excessive water on the
top. The amount of water added was gradually reduced to 5 ml for Runs K-IO through
K-14. The products from the last four runs were stored and decanted to use as the feed
solution for Run KC-2.
The same catalyst, KN-A, was tested again using the water added feed solution for
7 hours (Run KC-2). The nitrogen and sulfur contents of the hourly products are sum­
marized in Table VI. The effect of addition of water to the pretreated feed solution on the
catalytic activity is illustrated in terms of denitrogenation and desulfurization (Figure 11).
It appears that the effect of 2.5 wt % water present in the feed was insignificant. In fact, a
slight degradation in catalytic activity with the water could partially be attributed to the
lower feed reservoir temperature. A normal reservoir temperature of IOO0C had to be
37
Table VI
Effect of Water Addition on Catalytic Performance
Feed:
SRL d i s s o l v e d i n
C ataly st:
KN-A w i t h
Run c o n d i t i o n s
N itrogen
Hour
0.5
1 . 0
1 .5
2 . 0
2.5
3.0
3.5
4.0
4.5
5 .0
5.5
6 . 0
6.5
7.0
Feed
MoO3
NiO
CoO
WO3
: 11.0
:
2 .7
:
4.8
: 10.2
T em perature:
Pressure
:
LHSV
:
H z/feed
:
(1:1)
wt%
wt%
wt%
wt%
42 5±5°C
p sig
1000
1 .0
1 0 ,0 0 0
scf/b b l
& s u lfu r analyses
No w a t e r
( DeN%)
(94.7)
0 . 0 2
(94.7)
0 . 0 2
0 . 0 2
(94.7)
0 . 0 2
(94.7)
0 . 0 2
(94.7)
0 . 0 2
(94.7)
0 . 0 2
(94.7)
0 .0 4 (89.2)
0 .0 6 (83.8)
(73.0)
0 . 1 0
—
—
—
0 .3 7
N
te tra lin
added
S
(DeS%)
0.19
0.16
0 .14
0 .1 3
0.15
0.14
0 . 2 1
0.17
0.30
—
—
0.47
Y i e l d : 6 7 . 5 vo I . %
D r a i n e d volum e = 19.
(59.6)
(67.0)
(70.2)
(72.3)
(6 8 . 1 )
(71.3)
(56.4)
(63. 8 )
(36.2)
0
ml .
2.5% w a t e r a d d e d
( De N%)
S
(DeS%)
(94.7) 0 . 1 2 (68.4)
0 . 0 2
0 . 0 2
(94.7) 0 .0 9 5 ( 7 5 .0 )
0 . 0 2
(94.7) 0 . 1 2 (68.4)
0 . 0 2
(94.7) 0 .1 0 5 ( 7 2 .4 )
0 . 0 2
(94.7) 0.09 (76.3)
0 . 0 2
( 9 4 . 7) 0 . 1 5 ( 6 1 . 8 )
0 .0 3 (92.1) 0 .1 4 (64.5)
0 .0 5 ( 8 6 . 8 ) 0 ,1 9 (50.0)
0 .1 3 (65.8) 0 .2 3 (40 .8 )
0 .1 3 (65.8) 0 .2 6 (31.6)
0 .1 6 (57.9) 0 .0 8 (78.9)
0 . 2 2
(42.1) 0 .1 6 (57.9)
0 .2 5 (34.2) 0 .2 8 (27.6)
0 .2 6 (31.6) 0 .3 2 (27.6)
0 .38
0.38
N
7 9 . 2 VOI . %
2 3 . 3 ml
100
80
Z
OZ
H O
Eh H
Eh
<
N
S H
§
Eh k
H
2 P
cm
S
8
dP
U
OJ
OO
De N% w i t h o u t w a t e r
DeS%
A
A DeN% w i t h 2.5% w a t e r
A DeS %
3
4
HOURS ON STREAM
Figure 11. Effect of water addition in wt% denitrogenation and wt% desulfurization.
39
lowered to 80°C to prevent the boiling of excessive water still present. Lipsch and Schuit
(1969) observed strong adsorption of water on the surface of molybdenum-cobalt catalyst
supported on alumina, resulting in poisoning of HDS of thiophene and hydrogenation of
butenes [60]. This suggests that water might be adsorbed on active sites,leaving less avail­
able to the reactions. However, an increase in yield was experienced when water was added.
Two previous sets of autoclave operating conditions were arbitrarily adopted to pro­
vide the feed solutions for Run KC-I and the water-added feed solution for Run KC-2.
The autoclave operating conditions were reevaluated by a factorial experiment. Runs K-15
through K-27 were made for this purpose. The SRL dissolubility in tetralin was gauged by
the weight percent sediment of centrifugation. The pretreated feed solution was cooled to
the room temperature and centrifuged at 2,200 rpm for 5 minutes to determine the
weight percent sediment. Zero percent sediment means a complete dissolution. Any value
less than 10% was considered as “good.” Three factors in this experiment were considered
as rows, groups and columns, each at two, two, and three levels respectively. This gave 12
possible combinations. The results are summarized in TableVII and presented in Figure 12.
As can be seen from the figure, the effect of retention time on the dissolubility is most
significant.
The ranges of the three key variables (temperature, hydrogen pressure, and retention
time) were further extended for the final optimization. The results of thirteen autoclave
runs are summarized in Table VIII. Run K-28 was made to see if the extended retention
time (4 hours) would enhance the dissolution. Contrary to the expectation, a hard solid
deposit was observed on the wall, resulting in a record high of 62.3% sediment. The same
operating conditions were used for Run K-29 except a retention time of 0 hour. The result
was good with 6.9% sediment without any visible carbonaceous deposit.
The reaction temperature was reduced to 525°C from 575°C in Run K-30. The
weight percent sediment was increased to 17.7% for zero hour retention time. This value
40
Table VII
SRL Dissolubility at Various Operating Conditions
R etention P ressu re
tim e, h r
psig
Wt% s e d i m e n t
4750C
425°C
0
56.3
(K - 1 5 )
40.6
(K-16)
37.5
(K - 1 7 )
0
1000
46.3
(K -1 8 )
41.9
(K - 1 9 )
38.6
( K -2 0 )
I
0
51.4
(K -2 1 )
32.0
(K - 2 2 )
18.7
(K -2 4 )
I
1000
60.3
(K - 2 5 )
34.9
(K-26)
15.5
(K -2 7 )
0
*
3750C
5250C
44.7
( K -2 3 )
* A r e t e n t i o n tim e o f 0 h r means t h a t th e a u t o c l a v e
w as b r o u g h t t o t h e d e s i r e d t e m p e r a t u r e a n d t h e n
t h e h e a t e r was t u r n e d o f f im m e d i a t e l y . See b e lo w .
I hr reten tio n
0
hr reten tio n
HOURS IN REACTION
WEIGHT PERCENT SEDIMENT
70
_|____________I____________I <<------------- 1-------------------1----------------- 1—
375
425
475
375
425
475
REACTOR TEMPERATURE, °C
Figure 12. Effect of reactor temperature and retention time at two levels of hydrogen pressure on SRL dissolubility in tetralin.
42
Table VIII
SRL Dissolubility at Extended Operating Conditions
R eten tio n
tim e
hr
Hydrogen
pressure
p sig
0
1000
0
2000
1 /6
2000
1/3
2000
1/3
3000
1 /2
2000
I
1000
I
2000
2
1000
4
1000
4
2000
425°C
Wt% s e d i m e n t
475°C
5 2 5°C
43.6
(K-36).
35.0
(K - 3 7 )
27 .3
(K-39)
6 .7
(K -40 )
6.9
(K-38)
—•
46.4
(K -3 3 )
4 .7
(K-35) .
5.9
(K - 3 2 )
—
1 2 . 6
( K - 3 4)
17.7
( K -3 0 )
—
5 75°C
6 .9
(K - 2 9 )
-
-
-
-
-
-
-
-
-
3.5
(K -3 1)
' -
6 2.3
. ( K -2 8 )
—
'
43
was reduced to a record low of 3.5% when the retention time was increased to I hour as
experienced with Run K-31. However, due to frequent problems with the electrical heat­
ing unit the maximum allowable temperature had to be limited to 475° C. Therefore,
the remainder of the batch runs were carried out to optimize the other two variables,
hydrogen pressure and retention time. The weight percent sediments obtained from three
runs (K-29, K-30, and K-31) were added and the results of the extended experiments are
shown in Figure 13. The dissolution improved as the reaction temperature increased. The
two curves, zero hour retention vs. one hour retention, ran in parallel at a weight percent
sediment of less than 37%. This trend indicates that the reduced reaction temperature
could be compensated to some extent by longer retention time.
Next, attention was focused on the role of the initial hydrogen pressure in dissolution.
Figure 14 showed two contrasting curves. The lower curve was constructed based on the
data obtained under the various initial hydrogen pressures at a fixed retention time (I hour)
and at a constant reaction temperature (475° C). The upper curve was plotted with the data
obtained under the same operating conditions except a constant reaction temperature of
425 C. It appeared that the higher initial hydrogen pressure enhanced the dissolution of
SRL in tetralin only at a certain reaction temperature (475° C) or higher (Figure 14).
Finally, efforts were made to determine the optimum retention time with the other
two fixed variables (475°C,and 2,000 psig H2). Figure 15 indicated that there was a drastic
improvement in dissolution when the retention time had been increased from 20 minutes
to 30 minutes. It appeared that a retention time of at least 30 minutes should be allowed if
the reaction temperature of 475°C and the initial hydrogen pressure of 2,000 psig were
chosen. It was also found that the initial hydrogen pressure of 3,000 psig could reduce the
retention time from 30 minutes to 20 minutes. However, the initial hydrogen pressure of
2,000 psig was more practical with the existing equipment , at hand. The optimized oper­
ating conditions, for the maximum dissolution are:
Re t e n t i o n
WEIGHT PERCENT SEDIMENT
O
A
tim e
hour
I hour
0
P r e s s u r e : 1000
10 -
425
475
52
REACTOR TEMPERATURE, °C
Figure 13. Effect of reactor temperature on SRL dissolubility in tetralin at two levels of retention time.
45
R eten tio n
tim e: I h r
T em perature
(O) 42 5°C
(A) 475°C
O
1000
2000
INITIAL H2 PRESSURE, PSIG
Figure 14. Effect of hydrogen pressure on SRL dissolubility at two different temperature
levels.
T e m p e r a t u r e : 4 75°C
H2 p r e s s u r e :
(A) 2000 p s i g
(®) 3000 p s i g
0
20
30
40
50
RETENTION TIME, MIN.
Figure 15. Effect of retention time on SRL dissolubility in tetralin at AlS0C and 2000
psig of H2.
I
46
1. Reactor temperature = 475° C
2. Initial hydrogen pressure = 2,000 psig
3. Retention time = 40 minutes
The batch runs from K-40 through K-82 were made with these optimum operating
conditions for the common feed solution. The feedstock thus prepared was used to evaluate the 27 catalysts in the continuous trickle bed reactor.
Performance Tests
A total o f thirty-seven runs have been, completed with a continuous trickle bed reac­
tor; a randomized complete design of twenty-seven catalysts, fractional replications of 9
catalysts, and a blank carrier. The first two runs were carried out for 8 hours; the first with
blank carrier and the second with K-14 catalyst. The nitrogen contents of hourly product
samples are presented in Table IX. Performance of the KT-14 catalyst is illustrated in
Figure 16 in comparison with that of a blank carrier. Although the K-14 catalyst is super­
ior in denitrogenation to the blank carrier, there is a striking similarity between these two
curves. The abrupt changes in denitrogenation during the second and third hour on stream
were partially attributed to the gradual plugup of pores by the suspended solid particles
in the feed and partially to the wide pore distribution of the catalysts. The latter is
explained further in “Development of A Model for Catalyst.” The remainder of catalysts
were tested for four hours and the product samples were taken hourly for analysis. The
reactor operating conditions are kept constant at 425±5°C, 1,000 psig, LHSV of 1.0 and
hydrogen flow rate of 10,000 scf per barrel of feed.
The nitrogen contents of hourly products are shown in Table X. The average weight
percent nitrogen concentration was calculated based on the hourly nitrogen values and.
the amount of the hourly product. The average denitrogenation (Table X) was based on
47
Table IX
Catalytic Performances: KT-14 vs. Blank Carrier
Feed:
SRL d i s s o l v e d i n
Run c o n d i t i o n s
te tra lin
(1:1)
T em perature:
425±5°C
Pressure
:
1000 p s i g
LHSV
;
1.0
H2Z f e e d
: 10,000 s c f / b b l
C a t a l y s t : (A) K T - 1 4
B ase: K a t a l c o 81-6731
w i t h 10 wt% MoO3
4 wt% CoO
8 wt% WO3
(B)
N itrogen an aly se s
K T - 14
B lank
Hr ~N
( DeN%)
N ( DeN%)
T 0 .0 1 (9 7 .8 ) 0 .1 5 (6 7 .4 )
2 0 .0 3 (9 3 .5 ) 0 .1 8 (60.9)
3 0 .1 7 (6 3 .0 ) 0 .3 3 (2 8 .3 )
4 0 .2 0 (5 6 .5 ) 0 .3 7 (1 9 .6 )
Y ield f o r 8 hours
90.0 voI %
B l a n k c a r r i e r (KT-O)
K a ta lc o 81-6731
w i t h no m e t a l s
KT-1 4
Hr 1 5
( DeN%)
5 6 .2 2 (5 2 .2 )
6 0 .2 4 (4 7 .8 )
7 0 . 2 7 ( 4 1 . 3)
8 0 .2 9 (3 7 .0 )
vs.
B lank
~N
( DeN%)
0 . 39(15.2)
0 .4 1 (1 0 .9 )
0 .3 8 (1 7 .4 )
0 .3 7 (1 9 .6 )
7 9 . 2 vol%
ASTM D- 8 6 D i s t i l l a t i o n
C h a r g e v o l u m e : 50 ml o f
the f i r s t 4 h our
com posite p ro d u c t
T e m p e r a t u r e , Qp
V ol.
K T - 14
B lank
IBP
243
286
5
378
382
10
380
386
15
390
390
39 4
394
20
23.5
4 36
—
25
402
422
30
—
35
64 0
Therm ow ell
F e e d —*- ~
1/4" D enstone
175 ml —
1/ 8 " D e n s t o r i e
25 ml
M ixture o f
cataly st
60 ml
1 /8 " Denstorie
60 ml
1/8"
D enstone
35 m l ^ '
KT- 1 4 C a t a l y s
B lank C a r r i e r
DENIT ROGENATION
O
•
HOURS ON STREAM
Figure 16. Catalytic performance of KT-14 against blank base in denitrogenation.
49
Table X
Nitrogen Removal With KT Series Catalysts
C at.
No.
B lank
KT-I
K T -2
K T-3
KT-4
KT-5
KT- 6
KT-7
KT- 8
K T- 9
KT-IO
K T-Il
KT-1 2
K T -1 3
KT-1 4
KT-1 5
KT-16
KT-1 7
KT-18
K T - 19
K T - 20
KT-2 1
KT-2 2
KT-2 3
KT-2 4
K T -2 5
KT-2 6
K T -2 7
K1T - I R
KT-3R
KT-7R
KT-9R
KT- 14R
KT- 19R
K T-2 IR
KT- 25R
KT-27R
N itroqen c o n c e n tra tio n .
H ours on s t r e a m ■
I
2
3
4
0.1 5
0 .1 8
0.33
0 .37
0.05
0 . 1 1
0.2 4
0.25
0.04
0 . 0 1
0.2 3
0 . 2 0
0.03
0.08
0.24
0 .27
0.03
0.17
0 . 0 1
0 . 2 2
0 .0 3
0 . 0 1
0.16
0.23
0.03
0 . 0 1
0.23
0.16
0 .0 7
0 . 0 1
0 .1 8
0.25
0 . 0 1
0.04
0.23
0.16
0 . 0 2
. 0 .0 5
0 . 2 1
0 . 2 0
0 . 0 2
0.0 4
0.19
0.2 4
0 . 0 1
0.04
0 .17
0.2 4
0 . 0 2
0.05
0.2 4
0.25
0 . 0 1
0.05
0.14
0.19
0 .0 3
0 . 0 1
0 .1 7
0 . 2 0
0 . 0 2
0.0 5
0.1 4
0 .2 3
0.03
0 . 0 1
0.16
0.25
0.0 3
0 . 0 1
0.1 6
0 . 2 2
0 . 0 2
0.05
0 . 2 0
0 . 2 0
0 . 0 1
0.05
0.19
0.25
0 .0 7
0 . 0 1
0.24
0 . 2 0
0 . 0 1
0 . 2 1
0.06
0.2 4
0 .04
0.05
0.18
0.24
0 . 0 2
0.08
0 . 2 2
0.19
0.0 3
0 . 0 1
0 . 2 2
0.26
0.04
0 . 0 1
0.1 6
0.23
0 . 0 2
0.05
0.28
0.19
0 . 0 1
0 .2 3
0.0 5
0 .2 5
0 .0 4
0.24
0 . 1 0
0.25
0.03
0.07
0.25
0.28
0 . 0 2
0 .0 7
0 .2 3
0.26
0 .0 3
0.06
0 . 2 2
0 . 2 0
0 .0 3
0 . 0 1
0.16
0 . 2 0
0 . 0 2
0 . 2 1
0 . 0 6
0.26
0 . 0 1
0 . 0 6
0 . 2 2
0.26
0 . 0 2
0.05
0.18
0.2 3
0 . 0 2
0.05
0 .2 3
0.2 8
wt%
Avq
0 . 2 76
0.184
0.146
0.179
0.135
0.135
0.137
0.168
0.131
0.144
0.149
0.143
0.166
0.113
0.125
0,140
0.136
0.132
0.137
0.158
0.15 7
0.169
0 . 1 2 0
0.148
0.16 3
0 .138
0.15 7
0.170
0.182
0.183
0.176
0.149
0 . 1 2 2
0.162
0.166
0.144
0.176
De N%
40.0
59.9
6 8 .3
6 1 .1
70.7
70.7
70.2
63.4
71.5
6 8 . 8
67.6
68.9
64.0
75.4
72.8
69 .6
70.4
71. 3
70.2
65.6
65 .9
63.2
73.8
67.9
64.6
70.0
65.8
63.0
60.5
60.3
61.5
67.5
73.5
64.7
64.0
68.4
61.7
50
the average weight percent nitrogen content of the products and the initial nitrogen con­
tent of the feed (0.46%).
The result of the ASTM D-86 distillation for all the composite products representing
each catalyst and blank carrier are presented in Appendix C. From these data, typical
ASTM distillation curves wpre constructed to illustrate the stepwise transformations of
SRL to distillate fuel. Figure 17 shows the two starting materials (SRL + tetralin), pre­
treated feed solution, and the catalytically hydrocracked product with the KT-17. The
blank carrier showed some influence on hydrocracking of the SRL. The best results were
obtained with the K-17 catalyst (10% MoO3,6% CoO, and 8% WO3). The hydrocracking
parameter was obtained from the ASTM distillation curve. Due to difficulties in dealing
with the curves, two criteria (hydrocracking parameters) have been introduced: (I) the
percentage of gasoline distilled up to 380°F (193°C) and (2) the percentage of residue
remaining in the pot at 450° F (232° C). These two values were used to characterize the
distillation products. The lower the percentage of residue with a larger portion of gaso­
line, the better.
The average desulfurization obtained from the sulfur content of composite products
and the feed is also presented in Table XI along with the hydrocracking parameter (%
distillate) and product yield.
Effect of Metal Compositions on Upgrading
The effect of three active metals impregnated on Katalco carrier on upgrading of
SRL were evaluated in three major categories: denitrogenation, desulfurization, and
hydrocracking. Each of these dependent variables was correlated with three indepen­
dent variables (factors): MoO3, CoO and WO3. A statistical analysis was performed with
the aid of a computer to obtain'multiple linear regressions using data transformations [61].
THERMOMETER READING IN
51
A T etralin
# SRL
O 1 s t 4- h r c o m p o s i t e
p r o d u c t w i t h K T - 17
Q Feed
0
8- h r com posite
p ro d u c t w ith blan k
base
PERCENT RECOVERED
Figure 17. ASTM D-86 distillation, feed and 8-hr composite product of blank carrier.
52
Table XI
Summary of Catalytic Performance
C
a
t
No.
DeN %
B lank
40.0
KT-I
59.9
KT-2
6 8 .3
KT-3
6 1 .1
KT-4
70.7
KT-5
70.7
KT- 6
70.2
KT-7
63.4
KT- 8
71.5
KT-9
6 8 . 8
KT-IO
. 67.6
K T-Il
68.9
K T - 12
64.0
K T - 13
75.4
K T - 14
72.8
K T - 15
69.6
K T - 16
70.4
K T - 17
71. 3
K T - 18
70.2
K T - 19
65 .6
K T - 20
65.9
KT-2 1
63.2
KT-2 2
73.8
KT-2 3
67.9
KT-2 4
64.6
KT-2 5
70.0
KT-2 6
65 .8
KT-2 7
63.0
KT-IR
60.5
KT-3R
60.2
KT-7R
61.5
KT-9R
67.5
K T - 14R
73.5
K T - 19 R
6 4 .7
K T - 2 IR
64.0
K T - 2 SR
68.4
KT -2 7R ' 6 1 . 7
.
% D istilla te
DeS % 3 8 0 UF 45 0 ° F
4. 5
9
57
52.5
17
63
57.5
20
69
55.3
29
71
69.0
13
62
72.1
24
65
70.6
26
70
61. 5
15
62
70.5
23
62
69.0
72
30
6 8.7
71
30
67.4
25
62
65.6
29
71
76.5
26
72
77.7
62
20
72.7
25
62
66.9
72
30
72.5
34
72
70.5
30
72
58.5
72
30
56.5
10
65
54.6
20
70
70.0
68
10
65 .1
13
67
67.5
67
10
63.9
. 17
67
6 8 . 0
13
. 68
50.0
68
10
54.2
16
64
53.6
28
70
56 .8
17
62
68.5
28
. 72
77.2
21
63
51.8
30
72
56.3
19
70
65 .7
18
67
5 3 .7
12
70 ■
W
Y ield
5 3 .3
72.3
70.2
79.4
7 2 .1
72.3
70.4
79.8
70.2
72.9
74.8
73.5
73.8
72.7
71.7
74.2
71. 7
72.9
75.6
74.2
74.2
67.3
74.8
76.3
73,8
73. 8
76.3
74.2
78.8
75.8
75.8
77,9
75.6
75.8
79.0
73.8
75.8
53
Figure 18 was plotted to show the weight percent denitrogenation as a function of
weight percent MoO3 neglecting the presence of the other two metals. The role of these
active metals in denitrogenation was clearly seen as compared to the blank carrier, and
the quadratic effect of MoO3 concentration was significant. Further breakdown to three
levels of CoO and WO3 concentrations is illustrated in Figure 19. Each curve shows a quad­
ratic trend with a local maximum denitrogenation at around 10% MoO3 concentration and
there are some interactions. Further increase in MoO3 content failed to demonstrate any
benefit. The global maximum denitrogenation was attained approximately at 10% MoO3,
4% CoO and 6% WO3. The nitrogen removal as an explicit function of CoO is shown in
Figure 20 and the details of individual correlation is shown in Figure 21. These two figures
also confirmed that the maximum nitrogen removal was achieved at around 4% CoO con­
centration. The effect of tungsten concentration on denitrogenation is explicitly illus­
trated in Figure 22. Unlike the two previous curves the presence of 2% WO3 appeared to
be as effective as any higher concentrations. Figure 23 shows this trend in detail. At low
MoO3 level (4%), an increase in WO3 up to 8% generally improved the denitrogenation..
However, at higher MoO3 levels (10% and. 16%) any addition of WO3 beyond 2% level
actually impaired catalytic performance in denitrogenation.
An analysis of variance (ANOVA) for three factors at three levels is shown in. Table
XII. An unweighted means procedure was used for unequal replications. The ANOVA table
is a statistical technique used to determine those factors having a significant effect on deni­
trogenation. The first F value of the fifth column greater than 4.26 designates the factor
that is significant at a significance level of 5%. That is, if the parameter is determined to
be significant, there is only one chance in twenty of being in error. Since the F-values are
greater than 4.26, each parameter has a significant effect.
Table XIII shows multiple linear regression which was obtained by the use of back­
ward stepwise regression technique at a significant level of 5%. Thus, regression equation
WEIGHT PERCENT DENITROGENATION
54
WEIGHT PERCENT MoO3
Figure 18. Catalytic performance in denitrogenation as a function of MoO3 concentration.
CoO:
CoO: 4%
WEIGHT PERCENT DENITROGENATION
CoO: 2%
6
%
A 14%
16
0
WEIGHT PERCENT
16
0
MoO3
Figure 19. Catalytic performance in denitrogenation as a function of CoO, WO3, and MoO3 concentrations.
WEIGHT PERCENT DENITROGENATION
56
WEIGHT PERCENT
CoO
Figure 20. Catalytic performance in denitrogenation as a function of CoO concentration.
100
MoO3 : 16%
MoO3 : 10%
MoO3 : 4%
WEIGHT PERCENT DENITROGENATION
90
80
-
-
-
70
-
60 -
i
-
50
-
-
40
-
-
-
O 2% WO3
□
8 % WO3
A 14% WO3
30
20
i
-
:
I
0
-
—
2
I
4
I
I
6
0
2
I
4
:
I
6
I
0
2
I
4
I
6
WEIGHT PERCENT CoO
Figure 21. Catalytic performance in denitrogenation as a function of MoO3, WO3, and CoO concentrations.
WEIGHT PERCENT DENITROGENATION
58
WEIGHT PERCENT WO3
Figure 22. Catalytic performance in denitrogenation as a function OfWO3 concentration.
100
WEIGHT PERCENT DENITROGENATION
90 80 70 60 50 40 30 -
0 2 % CoO
□
4% CoO
A 6 % CoO
20 j:
j
___
I_ _ _ _ _ _ _ I_ _ _ _ _ _ _ _ _ _ _ _
O 2
8
i
14
I
I_ _ _ _ _ _ _ _ _ _ _ I_ _ _ _ _ _ _ _ _ _ _ _ I
02
8
14
T
i
i
_ _ _ _ _ _ _ _ _ _ I_ _ _ _ _ _ _ _ _ _ _ j_ _ _
0 2
8
14
WEIGHT PERCENT WO3
Figure 23. Catalytic performance in denitrogenation as a function of MoO3, CoO, and WO3 concentrations.
60
Table XII
Analysis of Variance for Three Factors at Three Levels
2s S t e b s k4II s tiIeIomii ofACELLSH & :
FACTOR
TRT
4
10
16
FACTOR
TRT
2
4
6
FACTOR
TRT
2
8
14
HARMONIC
S
N
9
9
9
=
N
9
9
9
=
N
9
9
9
CT=
I
MEANS
66.98
70.06
66.48
2
MEANS
64,92
70.67
67.93
3
MEANS
68.32
69.27
65.93
1,200
ANALYSIS OF IVARIANCE:
DF
S .S ,
SOURCE
2
I 2
3
I 3
2 3
I 2 3
RESIDUAL
I4
2
4
4
8
9
MEANS,
PER
14^9
13,96
53,46
69.47
36.92
23.96
5.275
F a c t o r I = Wt% MoO 3
F a c t o r 2 = Wt% CoO
F a c t o r 3 = Wt% WO3
H.S.
8:8
3.491
26.73
17,37
9.230
2,995
.5861
F-VALUE
P-VALUE
WA
5.955
o ! oo8
45.61
29.63
15.75
5.110
T h re e L e v e ls
4%
1 0 %
16%
2%
4%
6 %
8%
14%
2%
R e s p o n s e V a r i a b l e = Wt% d e n i t r o g e n a t i o n
,1292E
0.000
0.000
.7309E
. 1295E
61
Table XIII
Multiple Regression Analysis for Denitrogenation
BEPEtlBEtlT VARIABLE
= 4
INDEPENDENT VARIABLES= 1 2 3. 12 13. 14 17
FITl
VAR R-PART
I ,6615
i
m
12 - .5 9 4 9
13 -.6851
INTERCEPT
R-SQUARED
ANALYSIS
SOURCE
I
I?
--'311
=
-
37.43
.7705
OF VARIANCE!
DF
S .S .
2I
35
R
TOTAL
627.1
B
2.330
■
-.9410E -01
-1 .0 7 6
=
=
=
=
=
=
=
i
M
T
4.667
1 :2 3 0
♦2403E-01 -3 ,9 1 6
.2163
-4 .9 7 6
-m i
H.S.
69.02
5.140
4 = Wt% d e n i t r o g e n a t i o n
1
2
3
12
13
14
17
SE(B)
Wt% MoO3
Wt% CoO
Wt% WO 3
(Wt% MoO3 ) (Wt% MoO3 )
(Wt% CoO) (Wt% CoO)
(Wt% WO3 ) (Wt% WO3 )
(Wt% MoO3 ) (Wt% WO3 )
F-VALUE
13.43
P-VALUE
0.000
P-VALUE
0.000
0.000
♦3002E-02
0.000
0.000
62
(the weight percent denitrogenation regressed on the metal oxide concentrations) is con­
structed as follows.
Wt% DeN = 37.43 + 2.33 (MoO3) + 9.37 (CdO) + 1.35 (WO3)
- 0.094 (MoO3)2 - 1.076 (CoO)2 - 0.061 (WO3)2
- 0.053 (MoO3) (WO3)
The presence of MoO3, CoO, and WO3 significantly affected the denitrogenation and the
coefficients of quadratic and the interaction of the metal oxide concentrations were nega­
tive. The intercept at 37.43 was close to a denitrogenation value obtained with the blank
carrier.
The weight percent desulfurization is plotted against the weight percent MoO3 in
Figure 24 ignoring the presence of CoO and WO3. The catalytic abilities of these same
metals were also demonstrated in desulfurization. The correlations between the weight per­
cent desulfurization and MoO3 concentration are illustrated in Figure 25 at three different
levels of CoO. Within each of these CoO levels, three levels of WO3 concentrations were
also considered. It is clearly seen from the figure that there is an optimum MoO3 concen­
tration (around 10%) in each group of curves. The effect of weight percent CoO is
extremely significant when its concentration was increased from 2% to 4% but further
increase beyond 4% actually deteriorated catalytic desulfurization activities. The effect of
tungsten concentration in desulfurization was not so significant at all levels of MoO3 and
CoO. The desulfurization as a function of CoO is shown in Figure 26 and Figure 27. The
cobalt exhibited a most distinctive enhancement effect in desulfurization with peaks
around 4%. This trend was clearly visible at all three levels of MoO3 concentrations. The
sulfur removal as a function of WO3 concentration is shown in Figure 28. The catalytic
performance in desulfurization as a function of three metals are shown in Figure 29. The
figure indicates that there are significant interactions.
WEIGHT PERCENT DESULFURIZATION
63
WEIGHT PERCENT MoO3
Figure 24. Catalytic performance in desulfurization as a function of MoO3 concentration.
WEIGHT PERCENT DESULFURIZATION
100
£
O
□
A
20
-
S
IS
::
1
0
2% WO3
8% WO3
14% WO3
4
10
16
0
4
10
16
0
4
10
16
WEIGHT PERCENT MoO3
Figure 25. Catalytic performance in desulfurization as a function of CoO, WO3, and MoO3 concentrations.
O OO
WEIGHT PERCENT DESULFURIZATION
65
Figure 26. Catalytic performance in desulfurization as a function of CoO concentration.
100
MoO3 :
MoO3 : 10%
4%
MoO3 : 16%
WEIGHT PERCENT DESULFURIZATION
90
80
70
60
n
A
50
40
O
D
A
30
20
S
______I______ I______ I_____
0
2
4
6
0
I
:
I
2
2% WO3
% WO1
% WO3
I
4
I
6
i
0
1
2
1
4
1
6
WEIGHT PERCENT CoO
Figure 27. Catalytic performance in desulfurization as a function of MoO3, WO3, and CoO concentrations.
WEIGHT PERCENT DESULFURIZATION
67
0
2
8
14
WEIGHT PERCENT WO3
Figure 28. Catalytic performance in desulfurization as a function of WO3 concentration.
WEIGHT PERCENT DESULFURIZATION
50
-
40
O
□
A
30
2% CoO
4% CoO
6 % CoO
20
i__I
0 2
■
14
f
. .
0 2
.
8
■
14
02
8
14
WEIGHT PERCENT WO3
Figure 29. Catalytic performance in desulfurization as a function of MoO3, CoO, and WO3 concentrations.
69
An analysis of variance was made for data in the configuration o f a response variable
(desulfurization) with three classifying factors arranged in randomized complete block
(Table XIV). An unweighed means procedure was also used here for unequal replications.
Based on the F-value of the fifth column, the interaction between MoO3 and CoO is the
least significant. The use of backward stepwise approach produced a table of multiple
linear regressions at a significance level of 5% (Table XV). The regression equation thus
obtained is
Wt% DeS = 5.2 + 4.99 (MoO3) + 19.14 (CoO) + 2.03 (WO3)
- 0.23 (MoO3)2 - 2.19 (CoO)2 - 0.086 (WO3)2
- 0.071 (MoO3)(WO3)
The effect of the three metals on desulfurization was extremely significant. The quad­
ratic and interaction effects of these metals were negative as in denitrogenation. The inter­
cept at 5.2% was much the same as the desulfurization value of 4.5% obtained with the
blank carrier.
The degree of hydrocracking of SRL, that is, upgrading of hydrogen content is ex­
pressed in terms of two parameters: the percentage of gasoline distilled up to 380° F, or
193° C (light liquid) and the percentage of residue remained in the pot at 450° F, or 232° C
(heavy liquid). The volume percentages of light and heavy distillates are plotted as a func­
tion of MoO3 concentration, as a function of CoO concentration, and as a function of
WO3 concentration in Figure 30, Figure 3 1, and Figure 32.
Although there are significant interactions, it is generally seen that the maximum
amount of gasoline was produced at around 9% MoO3 concentration. The effect of CoO
concentration on gasoline production appears to be insignificant. The tungsten concen­
tration, on the other hand, played a significant role when the MoO3 concentration was
as low as 4%. In this case, at all levels of CoO concentration the gasoline yield increase
was almost proportional to the WO3 concentration (Figure 32). At higher MoO3 concen-
70
Table XIV
Analysis of Variance for Desulfurization
as Response Variable________
S
N
9
9
9
=
N
9
9
9
2
MEANS
59,34
71,22
65,88
FACTOR
TRT
N
2
9
8
9
14
9
HARMONIC CT=
3
MEANS
64,84
67.45,
64.16
1,200
FACTOR
TRT
4
10
16
FACTOR
TRT
2
4
6
I
MEANS
63,93
70,92
61.59
ANALYSIS OF VARIANCE:
SOURCE
DF
S»S
I
2 423.4
2
2 636,5
I 2
4 62.02
3
2 54,41
I 3
4 67.49
, 2 3
4 45,04
I 2 3
8 114.6
RESIDUAL
9 38.78
F a c t o r I = Wt% MoO 3
F a c t o r 2 = Wt% CoO
F a c t o r 3 = wt% WO3
M.S.
211.7
318.2
15.50
27.21
16.87
11.26
14,33
4.309
F-VALUE
49.13
73.85
3.598
6.314
3.916
2,613
3.325
Three L e v e ls
4%
10 %
16%
2 %
4%
6 %
2%
8%
14%
R e s p o n s e V a r i a b l e = Wt% d e s u l f u r i z a t i o n
P-VALUE
0.000
0.000
.5112E-01
. 1923E-01
,4125E-01
.1064
♦4622E-01
71
Table XV
Multiple Regression Analysis for Desulfurization
-•
DEPENDENT VARIABLE
'
-
‘ ----------
5 „
INDEPENDENT VARIABLES= I 2 3 12 13 14 17
F IT :
VAR
1
2
3
R-PART
.7 6 3 0
.7 9 1 8
.4 9 9 9
14 - .3 9 0 3
17 - .5 0 9 3
INTERCEPT
R-SQUARED
=
=
-♦ 8 6 2 6 E -0 1
- .7 0 5 6 E -0 1
iiii?
M. S .
2 6 7 .4
1 3 .1 6
5 = Wt% d e s u l f u r i z a t i o n
1
2
3
12
13
14
17
=
=
=
=
=
=
=
5 « )
2 .7 9 1
»6658
.3 8 4 5 E -0 1
,3 4 6 1
♦3845E-01
,2 2 5 3 E -0 1
T
6 .2 4 7
6 .8 5 9
3 ,0 5 4
- 6 .0 0 0
- 6 .3 2 5
- 2 .2 4 3
- 3 .1 3 1
5 .1 9 9
,8 3 5 5
ANALYSIS OF VARIANCE:
SOURCE
DF
S .S .
REGRESS
7 1 8 72.
? i ? i r L •I i
B
4 .9 8 9
1 9 .1 4
2 .0 3 4
Wt% MoO3
Wt% CoO
Wt% WO3
(Wt% MoO3 ) (Wt% MoO3 )
(Wt% CoO) (Wt% CoO)
(Wt% WO3 ) (Wt% WO3 )
(Wt% MoO3 ) (Wt% WO3 )
F-VALUE
2 0 .3 2
P-VALUE
0 .0 0 0
?4908E -02
.3 2 9 8 E -0 1
♦4051E-02
CoO: 2%
CoO: 4%
CoO:
6
%
Q
a
a
H
Eh
CO
450°F
H
Q
Eh
2
8
w
a
A
2% WO3
8 % WO 3
14% WO3
380°F
WEIGHT PERCENT MoO3
Figure 30. Catalytic performance in hydrocracking as a function of CoO, WO3, and MoO3 concentrations.
80
-
70
-
60
-
50
-
40
-
30
-
20
-
10
-
MoO3 : 4%
MoO3 : 10%
MoO3 : 16%
O
2% WO3
D
8 % WO 3
A 14% WO3
380°F
0
0
WEIGHT PERCENT CoO
Figure 3 1. Catalytic performance in hydrocracking as a function of MoO3, WO3, and CoO concentrations.
MoO3 : 16%
VOLUME PERCENT DISTILLED
MoOo: 4%
45 O0 F
45 O0 F
2% CoO
4% CoO
6 % CoO
38 0 o F
I
I____________ I____________ I__
0 2
8
14
WEIGHT PERCENT WO3
Figure 32. Catalytic performance in hydrocracking as a function of MoO3, CoO, and WO3 concentrations.
75
trations, this trend did not hold. An analysis of variance and a table of multiple linear
regressions made for the response variable, volume percentage of gasoline produced, is
shown in Table XVI and Table XVII.
Development of A Model for Catalyst
The Katalco earner (serial no. 81-6731) was inspected under the scan n in g electron
microscope. Figure 33 shows the scanning electron photomicrograph of catalyst surface
at 50,000 times magnification.' One millimeter in length in these figures is equivalent to
200A ( I A = I ten-billionth of a meter). A base granule as small as about 80A can be visu­
ally identified from the pictures. A summary of the granule size calculations is presented in
Appendix D. The calculations show that the average size of a base granule is 74A in diame­
ter. Figure 34 shows one of the many possible models that can explain the physical proper­
ties of the catalyst carrier. It shows a highly organized pore structure with an average pore
diameter of approximately 15 OA in three dimensional coordinates. In reality, however, the
pores are not well-defined cylindrical pores as evidenced by Figure 33.
Problems have been encountered in correlating the catalytic performance with a nom­
inal pore diameter and surface area. A number of commercial catalysts, well-known for ^
their denitrogenation capability with one type of feed, displayed poor performance with
another type of feed [63,64]. It is believed that the nominal pore diameter as used in the
monodisperse model is not an adequate parameter to represent a catalyst,particularly when
the large molecular substance such as coal-derived liquid was treated.
Stenberg [62] reported that the SRL had an average molecular weight of 460 with a
range of 160 to 4000 (gel permeation chromatographic separation). The SRL is highly
aromatic with one acid group and 0 .1 base group per molecule. The average thickness of
the aromatic layers in the SRL samples is about 12A, a stack of 4 planes with 15A in diam­
eter (powder X-ray diffraction).
76
Table XVI
Multiple Regression Analysis for Gasoline Yield
DEPENDENT VARIABLE
= 6
INDEPENDENT VARIABLES= I 2 12 13
FIT?
VAR R-PART
1 .5 2 4 5
2 - .3 9 5 9
12 - .5 7 1 4
13 .3 7 9 2
INTERCEPT
R-SGUARED
=
=
1
2
12
13
SE(B)
T
1 .2 4 8
3 .4 3 0
4 .4 7 6
- 2 .4 0 0
.6 1 6 4 E -0 1 - 3 .8 7 7
.5 5 4 8
2 .2 8 2
2 7 .6 8
.4 5 2 9
ANALYSIS OF VARIANCE?
SOURCE
DF
S .S .
REGRESS
4 8 8 6 .6
IBUAL
AL
19581
6
B
4 .2 7 9
- 1 0 .7 4
- .2 3 9 0
1 .2 6 6
M.S,
2 2 1 .7
3 4 .5 5
F-VALUE
6 .4 1 6
= Volume % g a s o l i n e y i e l d
— Wt% MoO3
= Wt% CoO
= (Wt% MoO3 ) (Wt% MoO3 )
= (Wt% CoO) (Wt% CoO)
P-VALUE
.9 5 0 4 E -0 3
.2 2 5 8 E -0 1
0.000
.2 9 5 SE-01
77
Table XVII
Multiple Regression Analysis for Heavy Oil Yield
BEFEfclSENT VARIABLE
=f>7.
INDEPENDENT VARIABLES= I 2 13 14 17
F IT .
VAR
I
2
13
14
17
INTERCEPT
R-SQUARED
=
=
R-PART
- .6 0 7 6
3443
- .3 3 9 0
- .6 4 7 4 ,
.5 8 6 0
♦
B
- .7 1 0 5
4 .4 9 4
- .5 4 7 4
t .5602E-<
.6 8 7 8 E -(
30
35
Il : 1W
H i:!
529.0
F-VALUE
5 .9 5 5
m
7 = Volume % h e a v y o i l y i e l d
1
2
13
14
17
=
=
=
=
=
T
- 4 ,1 9 0
2 .0 0 8
- 1 .9 7 4
1 :$I?
31»48
.4981
ANALYSIS OF VARIANCE:
DF
SOURCE
S »S *
EmIjAL
TOTAL
SE(B)
.1 6 9 6
2 .2 3 8
.2 7 7 3
Wt% MoO3
Wt% CoO
(Wt% CoO) **2
(Wt% WO3 ) **2
(Wt% MoO3 ) (Wt%W0
3
)
P-VALUE
.8484E
P-VALUE
0,00 0
.5 3 6 8 E -0 1
,5 7 6 9 E - 0 1
0.000
0.000
(A)
(B)
Figure 33. Scanning electron photomicrographs of various catalysts:
(A) Katalco blank carrier
(B) KT-14 with 10% MoO3,4% CoO, and 8% WO3 on Katalco carrier
(C) KT-14 after 8-hr run
(C)
(D)
(E)
Figure 33. (cont.)
(D) Union Carbide Linde 13X with 10% MoO3,4% CoO, and 8% WO3
(E) Ketjen LA-3P with 10% MoO3,4% CoO, and 8% WO3
(F)
Figure 33. (cont.)
(F) Nalco A with 10% MoO3,4% CoO, and 8% WO3
(G) Harshaw CoMo 0401 with 9% MoO3 and 3% CoO
(G)
81
Figure 34. Three dimensional arrangement of basic granules for Katalco carrier.
82
Since cellulose was believed to be one of the possible chemical precursors of coal
[14], nitrocellulose was chosen as the solute dissolved in the dibutyl phthalate. The result*
ing solution containing 0.35% nitrogen was used as the feed to evaluate five catalysts of
different nominal pore diameter. Nitrocellulose, or cellulose trinitrate (C6H7O2 (ONO2 )2)n,
was obtained from Hercules, Inc. The average number of anhydroglucose units in the
molecule ranged from 500 to 2,500 when chemically purified [65]. From this information,
the polymer size was estimated to be 15A in diameter and from 2,250A to 11,250A in
length. The five catalysts selected are Linde 13X molecular sieve of Union Carbide (nomi­
nal pore diameter (P.D. = IOA), LA-3P of Ketjen (P.D. =55A), 78-6008 o f Nalco A (P.D. 89A), MoCo 0401 of Harshaw (P.D. = 100A), and 81-6731 of Katalco (P.D. = 169A). Each
of these catalysts has 10% MoO3,4% CoO1 and 8% WO3 except the Harshaw having 9%
MoO3 and 3% CoO. They were tested in the continuous trickle bed reactor. The conditions
were:
1. Reactor temperature
350°C
2. H2 Pressure
I 1QOOpsig
3. LHSV
1.0
4. H2Zfeed
10,000 scf/bbl
The results of the 8-hour tests are presented in Table XVIII. The denitrogenation is plotted
against hours on stream for each catalyst in Figure 35. The Union Carbide 13X molecular
sieve having a nominal pore diameter of IOA deactivated rapidly in a few hours. The other
four catalysts performed in the order of their pore diameters except that the Nalco A
(P.D. = 89A) outperformed Katalco (P.D. = 169A). The pore distributions for both Nalco A and Katalco carrier were obtained from Katalcd Research Laboratory of Chicago
(Appendix A). From these data the actual surface area covered by the 10% MoO3 was cal­
culated for Katalco in Table XIX and for Nalco A in Table XX. The key assumption used
here is that the metal salts initially absorbed in the pores remained on the wall of the cor-
Table XVIII
Comparison of Various Catalysts
U nion
C arbide
13X
C arrier:
M aterial
P o r e v o lu m e , ml/gm
P o r e d i a m e t e r , °A
S u r f a c e a r e a , m2/ g m
G ra n u le s i z e , O A (dia.)
10
Wt%DeN o f p r o d u c t s :
9 3 .3
Feed:
N itro cellu lo se
(0.35
K atalco
N alco A
Harshaw
LA- 3P
81-6731
78-6008
CoMoO4 0 1
gamma
alum ina
0 .933
169
207
74
gamma
a lum ina
0.72
89
337
46
gamma
alum ina
0.4
10
10
M o l e c u l - gamma
a r s ie v e alum ina
0 .6 8
0.654
55
10
( 2 6 0 0 ?) 410
38
(6
?)
M etals
MoO 3 , wt%
i m p r e g n a t e d : CoO, wt%
WO3 , wt%
I hr
2 hr
3 hr
4 hr
8 hr
K etjen
10
4
4
8
8
-
99.5
33.3
9 6.7
25.7
6 4.3
n %)
4
9
3
8
8
0
99.5
.
96.2
-
*-
160
96
4
—
-
100
65.2
99.0
99.5
9 8.1
—
91.9
9 1 .1
r^ ^ v C O O C H 2 CH 2 CH 2 CH 3
< ^ > C O O C H 2C H 2C H 2 CH3
ONO2 Is CH 2 ONO 2
500 t o
2500
W
99.5
d isso lv ed in d ib u ty l p h th a la te
CH2 ONO2
00
N alco A
( 89°A)
H arshaw
( IOO0 A)
K atalco
(1 69° A )
K etjen
( 55°A)
Union
C arbide
( 10°A)
0
1
2
3
4
5
6
7
HOURS ON STREAM
S o l u t e : N i t r o c e l l u l o s e (0 .3 5% i n N)
S o lv en t: D ibutyl p h th a la te
Figure 35. Effect of catalyst pore diameter on denitrogenation.
8
9
10
Table XIX
Pore Volume and Surface Area Following 10% MoO3,4% CoO,
_______ and 8% WO3 Impregnation on Katalco Carrier*
*
yu u i i v u u - 6 0 0 )
500 ( 6 0 0 - 4 0 0 )
300 ( 4 0 0 - 2 0 0 )
190 ( 2 0 0 - 1 8 0 )
I / 0 ( 180-160)
150 ( 1 6 0 - 1 4 0 )
130 ( 1 4 0 - 1 2 0 )
HO ( 1 2 0 - 1 0 0 )
90 ( 1 0 0 - 80)
70 (
8 0 — bO)
50 ( 6 0 - 40)
30 (
4 0 - 25)
T otal
*
P .V .:
S.
T.
S.M .:
%ASM:
P.V .
ml
gift
0.192
0.146
0.269
0.034
0.036
0.041
0.044
0 .033
0 .043
0.043
0.037
0.015
0.933
S .A . T.M . S .M.
m2
m2
qm OA
gm
9 .8 4.00
9 .8
1 2 . 1
2.76 1 2 . 1
38.5 1 .6 0 2 3 .7
7 .2 1.08
3.0
8.3 0 .9 9
3.2
1 1 .0
0.8 5
3 .6
13.5 0 .7 5
3.9
16.2 0 .4 7
2.9
19 .1 0 .5 1
3.8
24.8 0 .4 0
3.8
29 .8 0.28
3.2
16.9 0 . 2 0
1 .3
20 7 .
74.3
%ASM
%Mo 0 3
gm
1 8.8
34.8
6 4.3
6 8 .0
72.0
76.5
8 1 .3
85.0
89.6
9 4 .4
9 8 .4
1 0 0 .
p o r e volum e
A .: su rfa ce area
M .: t h i c k n e s s o f MoO3
s u r f a c e a r e a c o v e r e d b y MoO3
p e r c e n t cum ulative s u rfa c e
a r e a c o v e r e d b y MoO3
P.M .
ml
gm
0.1881
0.1 4 2 7
0 .2628
0.0 3 3 2
0 .0351
0 .0401
0.0430
0.0322
0.0420
0.0420
0.0362
0.0 1 4 7
0.912
T .C . P .C .
ml
OA
gm
1 .3 1 0.1868
0.80 0 .1 4 1 7
0.46 0.2 6 1 1
0 .3 1 0.0330
0 .2 9 0.0349
0.25 0.0398
0 .2 2
0.0427
0.1 4 0.0320
0.15 0.0417
0 .1 2
0.0 4 1 7
0 .0 8 0.0359
0.06 0 .0146
0.9059
T.W. P .W.
ml
OA
gm
2.35 0.1845
1.44 0.1400
0 .8 4 0 .2578
0 .5 7 0.0326
0 .5 2 0.0344
0 .4 5 0.0393
0.39 0.0422
0 . 2 4 0 . 0 316
0 .2 7 0.0412
0 . 2 1
0.0412
0 .1 5 0.0355
0 . 1 1
0.0144
0.8947
P . M . : p o r e v o l u m e a f t e r MoO3
T . C . : t h i c k n e s s o f CoO
P . C . : p o r e v o l u m e a f t e r CoO
T . W . : t h i c k n e s s o f WO3
P . W . : p o r e v o l u m e a f t e r WO3
Table XX
Pore Volume and Surface Area Following 10% MoO3,4% CoO,
_______ and 8% WO3 Impregnation on Nalco A Carrier*
*
900(1200-600)
500 ( 6 0 0 - 4 0 0 )
300 ( 4 0 0 - 2 0 0 )
190 ( 2 0 0 - 1 8 0 )
170 ( 1 8 0 - 1 6 0 )
150 ( 1 6 0 - 1 4 0 )
130 ( 1 4 0 - 1 2 0 )
HO ( 1 2 0 - 1 0 0 )
9 0 ( 1 0 0 - 80)
70 (
80— 60)
50 ( 6 0 - 40)
30 (
4 0 - 20)
T otal
*
P .V .
S .A.
m2
ml
gm
qm
0.007
0 .3
0.006
0.5
0.029
4.5
0 . 0 1 2
0 . 0 2 0
0 .037
0.067
0.107
0.160
0.159
0 . 1 1 2
0 .0 0 3
0.72
T.M.
OA
6 .9 1
3.55
1 .9 1
1.37
2 . 6
4 .7 1.26
9 .7 1.13
0 .9 4
2 1 .0
39. I 0 . 8 1
71.7 0 . 6 6
9 0 .7 0.52
89.2 0 .3 7
3.3 0 .2 7
337.
S.M.
m2
qm
0.30
0.50
3.3 1
1.37
2.28
4.22
7.59
12.18
18.20
18.14
12.69
0.3 4
81.12
P . V . : p o r e volum e
S.
A .: s u r f a c e a re a
T.
M .: t h i c k n e s s o f MoO3
S . M . : s u r f a c e a r e a c o v e r e d b y MoO3
%ASM: p e r c e n t c u m u l a t i v e s u r f a c e
a r e a c o v e r e d b y MoO3
%ASM
%Mo C>3
qm
0 .9 7
1.8 1
5.85
7.53
10.31
15.46
24.75
39.64
6 1 . 89
84.06
9 9 .5 8
1 0 0 .0
P. M .
ml
qm
0 .0068
0.0058
0.0281
0.0116
0.0194
0.0359
0.0650
0 . 1 0 38
0.1553
0.1543
0.1807
0.0029
0.6976
P .M .:
T .C .:
P .C .:
T .W .:
P.W .:
T .C . P . G.
ml
°A
qm
2 . 0 1
0.0067
1 .0 3 0.0058
0 .5 6 0.0297
0.4 0 0.0115
0 .3 7 0.0192
0 .3 3 0.0356
0 .2 7 0.0645
0 .2 4 0.1029
0 .1 9 0.1539
0 .1 5 0.1529
0 . 1 1
0 . 1 0 77
0 .0 8 0.0029
0.6915
T.W. P. W .
ml
OA
qm
3.62 0.0 0 6 6
1 .8 6
0 .0057
1 .0 0
0.0274
0.72 0.0114
0 .6 6
0.0189
0.59 0.0350
0 .5 0 0.0634
0.42 0.1013
0.35 0.1514
0 .2 7 0.1505
0.19 0 .1060
0.14 0.0028
0.6776
p o r e v o l u m e a f t e r M0 O 3
t h i c k n e s s o f CoO
p o r e v o l u m e a f t e r CoO
t h i c k n e s s o f WO3
p o r e v o l u m e a f t e r WO3
87
responding pores during the drying period. This means that the surrounding wall (surface ,
area) having a large pore volume will have a thicker layer of metal. A plot of surface area
against the pore diameter for Katalco and Nalco A is presented in Figure 36. The calcu­
lations indicate that Nalco A accommodated a larger MoO3 covered surface area (8 1m 2/
gm) than the Katalco (74 m2 /gm). This could be one contributing factor of Nalco’s super­
iority over Katalco. Similarly, the MoO3 covered surface areas were calculated when the
16% MoO3 was impregnated into both catalysts. Again the Nalcp A gave a larger MoO3
covered area (128 m2/gm) than the Katalco (HO m2 /gm). On the contrary the perform­
ance of the catalysts with 16% MoO3 was generally inferior to that of catalysts with 10%
MoO3. During the 16% MoO3 impregnation the catalysts thus prepared tended to stick
together, indicating a possible dissolution of alumina in the highly concentrated MoO3
solution. This might have adversely affected the micropore structures.
The pore volume was progressively reduced as the three metals were impregnated in .
sequence. The calculated values agreed well with the measured volume within ±5%. Figure
37 shows how effectively the MoO3 is used in Nalco A as compared to Katalco. That is,
the majority of MoO3 impregnated is in a range of 50A to 180A in pore diameter. On the
other hand, in Katalco almost two-thirds of impregnation is in the pore of larger than
300A in diameter. The MoO3 layer is thicker throughout the pore distribution in Nalco A
than in Katalco.
The nature of the interconnection of the pores of various sizes is the key factor of the
model. Four different pore structures are shown in Figure 38. The first three arrangements
(Models A, B, and C) are a tridisperse model and the last structure (Model D) is the monodisperse model. AU of these four structures display the same physical properties; the same
pore volume, the same pore diameter, and the same surface area. If the molybdenum salt
solution is prepared for a monolayer (2.6A) coverage of MoO3 in the monodisperse model,
SURFACE AREA, m2/g m CARRIER
D
0
K a ta lc o
N a lc o A
PORE DIAMETER, OA
Figure 36. Surface area distribution vs. pore diameter.
89
10 0
□
O
K a ta lc o
N a lc o A
90
70
60
50
40
30
20
10
0
300 200
100
PORE DIAMETER, °A
Figure 37. Distribution of MoO3 vs. pore diameter.
PERCENT CUMULATIVE SURFACE AREA
80
S u rfa ce
7
VO
O
-T
B
C
D
u
In terior
P ore
d ia m e t e r
OA
50
100
200
T o ta l
Pore
v o lu m e
oa3
39 2 , 6 9 9
1,570,796
6,283,185
8,246,680
S u rfa ce
area
OA2
31,416
5,890
62,832
23,562
125,664
249,364
Figure 38. Representative structures of pore.
IOO0 A
n o
2 (TL) (L) — 8, 2 4 6 , 6 8 0
or
(D) (TL) (L) = 249 , 3 6 4
D = 132°A
f o r L = 600 °A
91
the same solution will make three different thickness of layers in each pore size of the
tridisperse model. They are:
Monodisperse Model
Monolayer
P.D.
132A
I
_____Tridisperse Model
Monolayer
PD .
50A
0.33
IOOA
0.67
200A
1.33
Therefore, it can be concluded that the monolayer coverage of the active metals using the
incipient wetness method is not possible, particularly when a catalyst has a wide pore dis­
tribution.
The pore structure and interconnection play key roles in the catalytic performance.
When the catalyst is completely wet with the feed liquid, the hydrogen will have to diffuse
through the pores into the active site. It can be reasoned that the wide pore opening at the
catalyst surface (see Model A of Figure 38) will allow more hydrogen to diffuse. The large
pore volume will contain enough hydrogen to supply the demand of the reacting liquid on
the active sites. As the pore size is reduced, both the dissolved hydrogen available and the
surface area are reduced as well. Therefore, this pore arrangement will favor the balanced
hydrocracking and hydrogenation [36A ]. On the other hand, in the pore model B of
Figure 38 the active sites are gradually increased,but much less hydrogen is available, favor­
ing the formation of carbonaceous deposit. This remains to be proven.
92
SUMMARY AND CONCLUSIONS
Based on the results of the three parts of the experiments, the following conclusions
are evident.
1. The solvent refined lignite (SRL) can be pulverized and dissolved in tetralin by
hydrotreating in a Parr series 4000 pressure reactor at elevated temperature and
pressure. The resulting liquid thus pretreated was suitable for catalystic upgrading.
The most favorable operating conditions for the batch runs are a reactor temper­
ature of 525°C, an initial H2 pressure of 3000 psig and a retention time of 20
minutes. The problems with the heater and the maximum allowable pressure of
the reactor led to an alternative optimization: 475°C, 2000 psig O f H 2 and a reac­
tion time of 40 minutes.
2. Two active metals (molybdenum and tungsten) and a promoter (cobalt) were
chosen to impregnate the Katalco-alumina as a carrier. A randomized complete
block design of the catalysts with three metals, each at three levels, were carried
out using a continuous trickle bed reactor for performance evaluation. The multi­
ple linear regression analysis indicates that the metal compositions for the best
hydrodenitrogenatipn (HDN) of the above pretreated feed are 10.5% MoO3,4.3%
CoO, and 6.4% WO3. The metal concentrations for the best hydrodesulfurization
(HDS) are much the same as those for the HDN. They are 9.64% MoO3, 4.37%
CoO, and 7.82% WO3. The effect of tungsten concentration on the catalytic hydro­
cracking was insignificant. The gasoline yield was maximized by the 8.95% MoO3
and 4.24% CoO. A high concentration of MdO3 as evidenced by 16% MoO3
impregnation appeared to adversely alter the original pore structure.
93
The modeling of a carrier structure with the aid of scanning electron photomicro­
graph reveals that all the catalyst carriers tested are constituted of base granules, or
sphericles with a diameter of 38A to 96A. A tracer, nitrocellulose with a known
molecular diameter dissolved in dibutylphthalate was used to determine whether
the nominal pore diamter can be a representative parameter for the selection of a
proper catalyst carrier. The results of the five 8-hour tests indicate that the nomi­
nal pore diameter as used in the monodisperse model is not an adequate parameter
to predict the catalytic performance. The incipient wetness method is acceptable
for the monolayer coverage of active metals if the carrier has a narrow band of
pore distribution (Nalco A). Building a monolayer by the use of incipient wetness
method is impossible in a carrier with a wide, band of pore distribution (Katalco).
The pore structure in addition to the pore distribution also plays a key role in
catalytic function. A wide surface area range with pore diameters of about 4 to 10
.
times as great as the critical diameter (solute molecular diameter/pore diameter)
will enhance the performance.
94
RECOMMENDATION FOR FUTURE STUDY
The nature of the interconnection of the pores of variable sizes appears to be one of
the key factors governing the life of a catalyst. A catalyst having a narrow band of pore dis­
tribution (monodisperse model) against a combination of two catalysts arbitrarily creating
a didisperse model should be tested. These experiments will isolate the effect of these pore
arrangements on the life of the catalyst from the other influencing factors.
LITERATURE CITED
96
LITERATURE CITED
1. Mottley, C. M., “How Much Energy Do We Really Need?,” Presented in Fuels and
Energy from Renewable Resources (Symposium volume), Academic Press, Inc.
(1977).
2. Anderson, L. L. and Tillman, D. A., Synthetic Fuels from Coal, John Wiley & Sons
(1979).
■
3. Office of Fossil Energy, US ERDA, Under Contract No. E(49-l8)-2225, Energy From
Coal: A State-Of-The-Art Review, p. I-I (1976).
4. Sergius, F-., “The Transformation of Coal into Oil by Means of Hydrogenation,” Pre­
sented at Proc. Intern. Conf. Bituminous Coal, November 15-18 (1926).
,
'
.
'
■
.
5. Martin, A. E. et al., Fuel Chemistry—A Mid-Century Perspective, Columbus, Ohio:
Battelle Memorial Institute (1974).
6. Nowacki, P., ed., Lignite Technology, New. Jersey: Noyes Data Corporation, pp. 3-12.
7. Rice, R. C., “Activities in Low-Rank Coal in the Northern Great Plains,” Presented to
the Eleventh Biennial Lignite Symposium, San Antonio, Texas, June 15 (1981).
8. Fodor, R. J., “Great Plains Gasification Associates Project,” Presented at the Eleventh
Biennial Lignite Symposium, San Antonio, Texas, June 15 (1981).
9. US DOE, Office of Fossil Energy, Under Contract No. EX-76-C-01-1224, Project Lig­
nite: Final Technical Report, March 28 (1972).
10. Considine, D. M., ed., Energy Technology Handbook, New York, McGraw-Hill, p. I17(1977).
11. Technology and Uses of Lignite, Proceedings of a Symposium Sponsored by the US
ERDA and University of North Dakota, Grand Forks, ND, May 18-19 (1977).
12. Severson, D. E., Souby, A. M., and Baker, G. G., “Continuous Liquefaction of Lignite
in a Process Development Unit,” Division of Fuel Chemistry, ACS (1977).
13. Farcasiu, M., Mitchell, T. O., and Whitehurst, D. D., Chem. Tech. 7, 680 (1977).
14. Whitehurst, D. D., “A Primer on the Chemistry and Constitution of Coal,”
15. Chen, A. S., “Flavanoid Pigments in the Red Mangrove, Rhizophora Mangle L. of the
Florida Everglades and in the Peat Derived from It,” Ph.D. Thesis, The Pennsylvania
State University, Sept. (1971).
97
16. Pines, A., Gibby, M. G., and Waugh, J. S., J. Chem. Phys. 59, 569 (1973).
17. Wender, I., “Catalytic Synthesis of Chemical From Coal,” Catalysis Review Science
and Engineering 14(1): 101 (1976).
18. Wiser, W., Preprints Fuel Division, ACS Meeting, 20(2), 122 (1975).
19. Orr, W., “Sulfur,” Handbook of Geochemistry, Vol. II/4, Springer-Verlag, Berlin,
Heidelberg-New York.
20. Cottrel, T. L., The Strength of Chemical Bonds, Butterworths Scientific Publications,
London, England (1958).
21. Satterfield, C. N. and Cocchetto, J. F., “Pyridine Hydrogenation: An Equilibrium
Limitation on the Formation of Pyridine Intermediate,” AIChE J. 21, No. 6, p. 1107
(1976).
22. Furimsky, E. and Amberg, C. H., “The Catalytic Hydrodesulfurization of Thiophenes.
VIII: Benzothiophene and 2-3-Dehydrobenzothiophene,” Can. J. Chem. 54, No. 10,
1507(1976).
23. Furimsky, E., “Catalytic Removal of Sulfur, Nitrogen, and Oxygen from Heavy Gas
Oil,” AIChE J. 25(2), p. 306 (1979).
24. Rossi, W. J. and Hudachek, R. J., Chem. Engr. 73(6), 138 (1966).
25. Kozlowski, R. H., Mason, H. F., and Scott, J. W., Ind. Eng. Chem. Proc. Des. Dev.
1(4), 276(1962).
26. Bray, B. G. and Kozlowski, R. H., LfS Patent 3,420,768, January 7 (1969).
27. Beuther, H., Donaldson, R. E., and Henke, A. M., Ind. Eng. Chem. Prod. Res. Dev.
3(3), 174(1964).
28. Hydrocarbon Processing, 55(9), 121 (1976).
29. Cheadle, G. D., “Unicracking-JHC Process Extends Commercial Applications,” Oil
and Gas Journal, July (1966).
30. Bond, G. C., Heterogeneous Catalysis: Principles and Applications, Clarendon Press,
Oxford (1974).
31. Harshaw Catalysts for Industry: Catalog No. 500.3, The Harshaw Chemical Co.,
Cleveland, Ohio.
32. Smith, J. M., Chemical Engineering Kinetics, McGraw-Hill, New York, 3rd ed., p. 3 5 1
(1981).
33. Cusumano, J. A., Catalysis in Coal Conversion, Academic Press, New York (1974).
34. Hucknall, D. J., “Selective Oxydation of Hydrocarbon,” Academic Press, New York
(1974).
98
35 • v T o 7h % « E., “Characterization of Molybdena Catalyst,” Advances in Catalysis.
V Ole
36
4bt I }
/O /e
S t O)8* C L ’ Catalytic Processes and Proven Catalysts. Academic Press, New York
36A.Clarke, I. K., Chem. Review, Vol. 7, p. 291 (1975).
37. Qader, S. A., I. Inst. Petrol. (London) 59, Np. 568, 178 (1973).
38. Clough, H., Ind. Eng. Chem. 49. 673 (19571
39
fl9 5 9 )r’ H” Flinn’ R" A' ’ ^
McKinley’ Jl B ’ Ind. Eng. Chem.. Vol. 51, p. 1349
40. Popov, I . , Chem. Abstr. 69, 62489r (1968).
41. Bliznakov, G., Popov, T., and Klisurski, D., Chem. Abstr. 70, 114535d (1969).
42. Perezhigina, I. Y. et al., Chem. Abstr. 72, 134857 (1970).
43. Satterfield, C. N., “Trickle Bed Reactors,” AIChE J. 21(2), 208 (1975).
44. Ross, L. D., “Performance of Trickle Bed Reactors,” Chem. Eng. Prog. 61fl0). 77
(1965).
45. US DOE, Solvent Refined Lignite Process Development: Quarterly Technical Progress
Report No. l,FE-4189-7 (1977).
:
46. Wagner, M. C., Letter to L. Berg, Nov. 21 (1978).
47. Emmett, P. H., ed., Catalysis, Reinhold, Vol. I, Chap. I (1957).
48. Schuit, G. E. A. and Gates, B. C., “Chemistry and Engineering of Catalytic Hydro­
desulfurization,” AIChEJ. 19(8), 417 (1973).
49. Stetcher, P. G., ed., The Merck Index, 9th ed., Merck & Co., Rahway, New Jersey, p.
633.
50. Sax, N. I., Dangerous Properties of Industrial Materials. 2nd ed., Reinhold, p. 888
(1963).
51. Norton Denstone 57 Catalyst Bed Supports, Bulletin D-12, Norton Company Cal­
gary, Canada.
52. Runnion, K. N., Catalytic Hydrogenation o f Synthoil,” M.S. Thesis, Montana State
University, March (1977).
.53. Cole-Palmer Catalog (79-80), p. 143.
99
54. Instructions for Operation of Brooks Thermal Mass Flowmeter, Brooks Instrument
Division, Emerson Electric Co., Hatfield, PA (1975).
55. Lake, G. R. et al., “Effect of Digestion Temperature of Kjeldahl Analysis,” Analytical
Chemistry 23(11), p. 1634 (1951).
56. American Society for Testing and Materials, “Standard Method of Test for Total
Nitrogen in Organic Materials by Modified Kjeldahl Method,” Annual Book of ASTM
Standards, Part 30, Designation E258 (1974).
57. Peters, E. D., Rounds, G. C., and Agazzi, E. J., “Determination of Sulfur and Halo­
gens,” Analytical Chemistry, VoL 24, p. 710 (1952).
58. American Society for Testing and Materials, “Standard Method of Test for Sulfur in
Petroleum Oils (Quartz Tube Method),” 1974 Annual Book of ASTM Standards, Part
23, ASTM Designation D 1551.
59. American Society for Testing and Materials, “ Standard Method of Test for Distilla­
tion of Petroleum Products,” Annual Book o f ASTM Standards (1974), Part 23,
ASTM Designation D-86.
60. Lipsch, J. M. and Schuit, G., “The CoO-MoO3-AI2O3 Catalyst III: Catalytic Proper­
ties,” L_Catalysis 15, p. 179 (1969).
61. Stenberg, V. I. et al., “Chemistry of Lignite Liquefaction,” US DOE, Quarterly
Report, Jan-Mar 1978, FE-2211-9.
62. Kujawa, S. T., “Catalytic Hydrogenation of Coal Derived Liquids,” Ph.D. Thesis,
Montana State University, July (1978).
63. Hass, G. R., “Catalytic Hydrogenation of Solvent Refined Coal,” Ph.D. Thesis, Mon­
tana State University, July (1978).
64. Nitrocellulose: Chemical and Physical Properties, Hercules, Inc., Willington, Delaware
(1979).
100
APPENDICES
101
APPENDIX A
SURFACE AREA AND PORE DISTRIBUTION
OF CATALYST CARRIERS
102
KATALCO c a r r i e r
SAMPLE # 8 1 - 6 7 3 1
96508 LC137-1901 (17-27/
N
I
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
RELATIVE
PRESSURE
(PZPO)
0.078
0.150
0.210
0.376
0.465
0.559
0.635
0.704
0.758
0.806
0.845
0.874
0.899
0.918
0.934
0.951
0.956
0.964
0.971
0.976 ■
0.981
0.993
0.987
0.983
0.978
0.971
0.963
0.958
6.941
0.925
0.905
0.879
0.852
0.810
0.764
0.707
0.633
0.550
0.455
0.350
0.200
0.139
0.080
S
3i-Aua-ei
GASEOUS N2
LIQUID N2
VOL SORP-EIi VOL SORBED
(CC/G e STP)
(CC/G)
51i555
'58.384
63.289
76.235
84.407
95,529
108.035
124.644
143.523
168.052
198.921
230.695
267.274
308.454
352.523
419.158
444.258
477.748
526.385
560.762
'593.366
709.652
680.541
660.895
631.346
602.085
568.041
553.453
487.523
437.839
383.404
328.670
282.910
230.226
187.265
149.988
119.442
98.707
80.736
69.529
57.010
51.874
45.907
0.080
0.091
0.099
0,119
0,132
0. 149
0.168
0.194
0.224
0.262
0.310
0.359
0.416
0.481
0.549
0.653
0.692
0.744
0.820
0.874
0.924
1.106
1.060
1.030
0.984
0.933
0.885
0.862
0.760
0.682
0.597
0.512
0.441
0.359
0.292
0.234
0.186
0.154
0.126
0.108
0.089
0.031
0.072
T
4.429
4.892
5.217
6.097
6.618
7.251
7.876
8,583
9.288
10.098
10,966
11.799
12.767
13.755
14.803
16.416
17.059
18.221
19.653
20.814
22.694
32.400
25.688
23.522
21.393
19.722
18.003
17.319
15.393
14.157
13.064
11.978
11.155
10.176
9.374
8.612 7.860
7.189
6.558
5.954
5.168
4,828
4.444'
A
16.
20.
23.
32.
38.
47,
58.
71.
87.
109.
135.
165.
204.
251.
308.
413. .
461 .
556.
692.
816.
1050.
2988.
1508.
1165.
884.
699.
537. .
481.
344.
272.
218.
172.
142.
111.
90.
72.
57.
46.
37.
30.
22.
19.
16.
DVZDP
0.147
0.129
0,121
0.143
0.185
0.256
0.375
0.544
0.796
1.233
1.737
2.262
3.307
4.401
6.036
7.520
6.776
10.525
11.902
9.236
14.730
6.978
7.848
8.370
7.474
6.027
5.051
5.972
4,778
4.349
.3.230
2.680
1.750
1.446
1.014
0.649
0.390
0.295
0.165
0.131
0.131
0.157
103
X
700
ISOTHERM FOR 9 6 5 0 8
L C 1 3 7 - 1 9 0 I (1 7 2 7 )
S
500
ML GAS/CM CATALYST
+
200
♦
♦
4
♦.
100
4.
X
♦
4.
♦
I ............I ............. I ............ I ............. I .............I .............I .............I .............I .............I
0
5
RELATIVE PRESSURE
1.0
<P/PO>
104
9 6 5 0 4 LC1 3 7 - 1 9 o i
(1 7 -2 7 )
S
3 1 - A u g - 8.1
ADSORPTION PORE VOLUME DATA
22 PO IN T PV AND 3 PO IN T SA
PORE
DlAMETER
A
600
500
400
300
200
195
190
185
180
175
170
165
160
155
150
145'
140
135
130
125
120
115
HO
105
100
95
90
85
80
75
70
65
60
55
50
45
40
35
30
25
0.1920
0.0628
0.0828
0.1166
0.1531
0.0085
0.0085
0.0086
0.0086
0.0086
0.0087
O.0087
0.0093
0.0096
0.0108
0.0103
0.0105
0.0107
0.0109
0.0110
0.0111
0.0112
0.0112
0,0111
0.0110 .
0.0108
0.0106
0.0106
0.0107
0.0109
0.0110
0.0107
0.0106
0.0102
0.0095
0.0089
0.0084
0.0077
0.0067
0.0003
CUMULATIVE
PORE VOLUME
(CCZG)
0.192
0.255
0.338
0.454
0.607
■0.616
'6.624
0.633
0.641
0.650
0.659
0.667
0.677
O*686
0.697
0.707
0.718
0.729
0.740
0.751
0.762
0.773
0.784
0.795.
0.806 •
0.817
0.828
0.838
0.849
0.860
0.871
0.882
0.892
0.902
0.912
0.921
0.929
0,937
0.943
0.944
MEDIAN PD ON SA = 115.9
MEDIAN PD ON PV = 287.0
VOL
PERCENT
(CCZG)
20.3
6.7
8.8
12.4
16.2
■
*0.9
0.9
0.9
0.9
0.9
0.9
0.9
1.0
1.0
I. I
I.I
I.I
I. I
1.2
1.2
1.2
1.2
1.2
1.2
1.2
I. I
I.I
I. I
I.I
1.2
I.2
I.i
I.I
1.1
1.0
0.9
0.9'
6.8
0.7
0.0
20.3
6.7
8.8
12.4
16.2
3.6
3.7
4. 4
4.6
4.7
2.3
2.3
2.3
2.3
Ci
1200 600 500 400 300 200 195 190 185 180 175 170 165 160 155 150 145 140 135 130 125 120 115 HO 105 100 95 90 85 80 75 70 65 60 55 50 45
40 35 30 -
PORE
VOLUME
(CCZG)
SAT PV = 1.1056
4V/A = 169.2
CO
BET SA =
223.15
1200 PORE VOLUME = 0.9438
I.5
SURFACE
AREA
(SO MZG)
CUMULATIVE
SURFACE .
AREA
9.81
9.8
4.57
14.4
7.49
21.9
13.47
35.3
25.05 .
60.4
1.72
62.1
1.78
•63.9
1.83
65,7
1.88
67.6
• 1.94
69.5
2.01
.71.5
2.09
73.6
2.28
75,9
2.44
78.4
2.83
81.2
2.80
84.0
2.96
86.9
3.12
90.1
3.28
93.3
3.46
96.8
3.63
100.4 '
3.80
104.2
3.97
108.2
.4.13
112.4
4.29
116.6
4.44
121 .I
4.60
125.7
4.84
■ 130.5 •
5.21
135.7
5.64
141.4
6.09
147.5
6.31
153.8
6.78
160.5
7.10
167.6
7.22
174.9
7.49
182.4
7.92
190.3
8.18
'198.5
8.22
206.7
0.48
207.2
105
KATALCO
D I G I S ORB 2500
SAMPLE:
96508 LC137-1901
(17-27)
S TA:
2 EQTIME: 3 METHOD:
ADSORPTION PORE VOLUME DISTRIBUTION
RANGE PORE
DIAMETER, A
+fC?P-500
4500-450
+450-400
+400-350
+35C-300
+300-280
*280-260
*260-240
+240-220
+220-200
+200-180
+180-1 6.0
+160-150
+150-140
♦140-130
+130-120
+120-110
♦110-100
+100 -95
+95 -90
+90 -85
+85 -80
*80 -75
+75 -70
+70 -65
+65 -60
+60 -55
*55 -50
+50 -45
+45 -40
+40 -35
+35 -30
+30 -25
+25 -20
AVERAGE
DIA.,A
+550.0
+475.0
+425.0
+375.0
+325.0
+290.0
+270.0
+250.0
+230.0
+210.0
+190.0
+1^0.0
+155.0
+145.0
+135.0
+125.0
+115.0
+105.0
+97.5
+92. 5
+S'7.5
+82. 5
+7-7.5
+72. 5
+67.5
+62. 5
+57.5
+52. 5
+4^.5
+42.5
+37.5
+32.5
+27.5
+22. 5
PORE VOLUME
(CC/G)’
+0.064137
+0.036491
+0.045089
+0.05249)
+0.068451
+0.02540,9
+0.030333
+0.030862
+0.032065
+0.036473
+0.033072
+0.039692
+0.017239
+0.022395
+0.021957
+0.021440
+0.026052
+0.020905
+0.008558
+0.012811
♦0.011823
+0.010291
+0.012139
+0.012493
+0.010126
+0.012341
+0.01 I348
+0.010594
+0.0]0780
' +0.010261
+0.010033
+0.009941
+0.010374
+0.014483
CUMULATIVE
PORE VOLUME
+0.064137
+0.I00628
+0.148717
+0.201209
+0.269660
+0.295069
+0.325902
+0.356764
+0.388829
+0.425302
+0.455374
+0.498066
+0.515306
+0.537701
+0.5596 S'7
+0.5.3I097
+0.607149
+0.623054
+0.636613
+0.649424
+0.661247
+0.67 1=37 .
+0.683676
+0.696170
+0.706296
+0.718637
+0.729935
+0.740578
+0.7513 5 9
+0.761620
+0.77]653 ■
+0.731594
+0.791969
+0.806452
SURFACE AREA CUMVLATI V
(SO M/G)
SURFACE AR
+4.665
+3.073
+4.526
+5.599
+8.425
+3.50 5
+4.568
+4.935
+5.577
+6.947
+6.962
+9.339
+4.449
+6.178
+6.506
+6.861
+9.062
+7.964
+3.51 I
+5.540
+5.405
+4.989
+6.265
+6.893
+6.001
+7.898
+7.894
. +8.07]
+9.078
+9.657
+10.702 .
+12.236
+15.090
+25.743
+4•665
•+7.737
+12.264
+17.863
+26.287
+29.792 '
+34.360
+39.293
+44.874
+51.822
+58.784
+68. I23
+72.572
+78."
7^O
+85.256 .
+92. II7
+101.173
+109.142
+112.653
+118.193
+123.598
•
*
•
128.507
+134.853
+141."46
+147.746
+155.644
+163.539
+171.610
+180.688
+190.346
+201.045 •
+213.284
+228.374
+254.121
3
106
96508
LC1 3 7 - 1 9 0 1 ( 1 7 - 2 7 )
S
31—Aug—8 1
ADSORPTION PORE VOLUME DATA
(DELTA V/DELTA D) * 1 0 0 0
IO
0
20
DIAMETER
- I
I ------I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
20
•
•
40
•
•
•
•
•
♦
I
I
I
60
«
.
•
•
•
•
I
I
I
I
I
I
I
I
I
I
I
A
80
100
120
•
•
•
•
.
•
140
160
180
200
107
96508 LC 137-1901
(17-27)
S
3 1-A ug-81 '
DESORPTION PORE VOLUME DATA
22 POINT PV AND 3 POINT SA
BET SA =
223.15
1 2 0 0 F1ORE VOLUME = 1 . Q3 3 7
PORE
DIAMETER
A
1200
600
500
400
300
200
195
190
185
180
175
170
165
160
155
150
145
140
135
130
125
120
115
HO
105
100
95
90
85
80
75
70
65
60
55
—
-
PORE
VOLUME
(CC/ G)
600
0.1388
500
' 0.0415
400
0.0638
300
0.1156
200
0.1786
195 . 0 . 0 1 2 4
190
0.0114
185
0.0115
180
0.0116
175 ■ 0 . 0 1 1 8
170
0.0122
165 '
0.0135
160
0.0141
155
0.0156
150
0.0153
145
0.0158
140
0.0163
135
0.0167
130
0.0171
125
0.0166
120
0.0173
115
0.0175
1 10
0.0177
105
0.0193
100
0.0208
95
0.0208
90
0.0212
85
0.0217
80
0.0219
75
0.0219
70
0. 02V7
65
0.0210
60
0.0200
55
0.0183
50
0.0024
SAT F-V = 1 . 1 0 5 6
4V/ A = 1 8 5 . 3
CUMULATIVE
PORE VOLUME
(CC/ G)
0.139
0.180
0.244
0.360
0.538
0.551
0.562
0.574
0.585
0.597
0.609
0.623
0.637
0.652
0.668
0.684
0.700
0.717
0.734
0,750
0.768
0.785
. 0.803
0.822
0.843
0.864
0.885
0.907
0.928
0.950
0.972
0.993
1.013
1.031
1.034
VOL
PERCENT
(CC/G)
13.4
4.0
6.2
11.2
17.3
1.2
I.I
1.1
1.1
1.1
1.2
1.3
1.4
1.5
1.5
1.5
1.6
1.6
1.7
1.6
1.7
1.7
1.7
1.9
2.0
2.0
2.1
2.1
2.1
2.1
2.1
2.0
1.9
1.8.
0.2
13.4 '
4.0
6.2
11.2
17.3
4.5
5.0
6.1
6.5
7i3
4.1
4.2
4.2
4.0
2.0
MEDIAN PD ON SA = 115
MEDIAN PD ON PV = 2 1 0
SURFACE
AREA
(SO M/G)
CUMULATIVE
SURFACE
AREA
6.80
3.03
5.79
13.30
29.51
2.51 .
2.37
2.45
2.55
2.65
2.83
3.23
3.46
3.97
4.03
4.30
4.58
4.87
5.15
5.21
5.65
5.95
6.29
7.20
8.11
8.52
9.18
9.90
10.61
11.31
11.95
12.46
12.80
12.71
1.84
6.8
9.8
15.6
28.9
58.4
60.9
63.3
65.8
68.3
70.9
73.8
77.0
80.5
84.4
88.5
92.8 •
97.3
102.2
107.4
112.6
118.2
124.2
130.5
137.7
145.8
154.3
163.5
173.4
184.0
195.3
207.2
219.7
232.5
2 45.2
247.1
108
KATALCO
DIGISORB 2500
SAMPLE: 9 6 5 0 8 L C 1 3 7 - 1 9 0 1 ( 1 7 - 2 7 )
2 EQTIME: 3 MODE: 7
METHOD: 2
DESORPTION PORE VOLUME DISTRIBUTION
PANGE PORE
DIAMETER,A
AVERAGE
DIA* » A
PORE VOLUME
<CC/G>
,»608-SOB, +550.0
+0.043961
+500-450
+475.8
+0.029866
+458-400
+425.0
♦0.040548
+400-350
+375.0
+0.052316
+350-300 . +325.0
+0.058604
+300-280
+290.0
+0.028096
+280-260
+270.0
+0.031599
+260-240
+250.0
+0.036114
+240-220
+230.0
+0.04299 I
+220-200
+210.0
+0.04001 I
+200-180
+190.0 , '+0.048256
+170.0
+180-160
+0.055122
+160-150
+155.0
+0.02=655
+150-I40
+145.0
+0.033893
+140-130
+135.0
+0.029842
+130-120
+125.0
♦0.035352
+120-110
+115.0
+0.041524
+110-100
+105.0
+0.037706
+100 -95
+97.5
+0.018525
♦95 -90
+92. 5
+0,024814
•+0.019400
+90 -85
+87.5
+85 -80
+82. 5
+0.021159
+80 -75
+77.5
+0.024718
+75 -70
+72.5
+0.023338
+70 -65
+67.5
+0.019187 .
+65 -60
+62. 5
+0.023145
+60 -55
+57.5
+0.020690
+55 -50
+52. 5
+0.0166 I3
+50 -45
+47.5
+0.019198
+45 -40
+42.5
+0.017634
+37.5
+40 -35
+0.014300
+35 -30
+32.5
+0.0.07251
+30 -25
+27.5
+0.001933
+25 -20
+2 2 . 5
+0.002392
S TA:
CUMULATIVE
PORE VOLUME
+0.043961
+0. 07382. 6
+0 . I 14374
+0.166691
♦0.225294
+0.253390
+0.284989
+ 0 . 3 2 1 103
+0.364094
+0.404105
+0.452361
+0.507433
+0.536139
+0.570032
+0.599373'
+0.635225
+0.676749
+0.714455
+0.732980
♦0.757794
+0.777194
+0.798353
+0.823071
+0.846409
+0.865596
+0.888741
+0.909431
+0.926044
+0.945242
+0.962876
+ 0. 9 7 - 7 ] 76
+0.984427
+0.986361
*0.988752
SURFACE AREA CUMULATIVE
(SG M/G)
SURFACE AREA
+3.197
+2.515
+3.816
' +5.580
+7.213
+3.875
+4.681
+5.778
+7.477
+7.621
+10.159
+12.970
+7.395
+9.350
+8.842
+11.313
+14.443
+14.364
+ 7 . <=00
+10.731
+8.868
+10.259
+ I2.758,
+ I2 . 8 7 6
+ 11.370
+14.313
+14.393
+12.657
+16.167
+16.597
+15.253
+3.925
+2.812
+4.253
+3.197
+5.712
+9.528
+ 1 5 . I 09
+22.322
+ 26 . I 97
+30.878
+36.656
+44.133
+ 5 1 . 7 54
+61.913
+74.883
+22,278
+91.628
+100.470
+ 1I 1.783
+126.226
+140.590
+ 1 4 8 . I 90
+158.920
+167.789
+ I72.043
+190.805
+203.682
+015.052
+229.864
+244.252
+256.915
+273.082
+289.678
+304.932
+313.857
+316.669
+320.921
X
109
96508 LCl37-1 901 (17 — 27/
S
Sl-Aua-Sl
PORE VOLUME DATA
ADSORPTION
DESORPTION
BET SA = 223.1 SO M/G
INTERCEPT = 0.000127
SLOPE = 0.019367
CUMULATIVE SA
= 207.2
CUMULATIVE SA
MEDIAN PD ON SA
= 115.9
MEDIAN PD ON SA
'
= 115.5
MEDIAN PD ON PV
= 287.0
MEDIAN PD ON PV
= 210.2
SATURATION PV
DIA
PV 1200
PV 1000
PV 800
PV 600
PV 400
PV 300
PV 200
PV 100
=
=
=
=
=
=
=
=
=
O
CK
AVERAGE PD(4V/A> = 169.2
AVERAGE PD(4VZA) = 185.3
SATURATION PV
CCZG
DIA
CCZG
.9438
.9143
.8610
.7518
«6063
.4897
.3366
.1377
PV 1200
PV 1000
PV 800
PV .600
PV 400
PV 300
PV 200
PV 100
=1.0337
-I.0035
= ,9583
= .8949
= .7896
= .6741
= .4954
= .1908
PV 1200 - 200 PD = 0.607
PV 1200 - 100 PD = 0.806
= 247.1
=
1.106
PV 1200 - 200 PD = 0.538
PV 1200 - 100 PD = 0.843
PERC PV 100-200 A
-IiIAh ON. 600 TPV =
26.5 ( .199)
PERC PV 100-200 A
DIAM ON .600 TPV =
34.0 < .305)
PERC PV 200-400 A
DIAM ON 600 TPV =
35.9 < .270)
PERC PV 200-400 A
DIAM ON 600 TPV =
32.9 ( .294)
no
9 6 5 0 8 LC1 3 7 - 1 9 0 1 ( 1 7 - 2 7 )
S
3 I - A u g - 81
SORPTION PORE VOLUME DATA
(DELTA V/DELTA D) *1000
0
10
I----I
I
I
20
I
DIAMETER
-I
A
20
I
I
I
.
'
40
I
I +
I
I
I
I
I
I
I
-
•
.
+
,
4
,
4
4
4
4
4
4
4
4
4
•
,
•
•
.
«
I
I
I
I
I
I
I
I
+
+
+
+
+
+
♦
+
+
+
•
.
•
«
«
+
+
•
+
+
+
+
+
+
I
I
I
I
I
.+
60
80
100
120
140
160
180
200
Ill
9 6 5 0 8 L C 1 3 7 -1 9 0 1 S
3I-A u g -8I
(DELTA PORE VOLUME/DELTA LOG PORE DIAMETER)
0
.7 0
7
I
.4 0
.4 0
.80
1.00
1.20
----------- 1 — ................I ------------- — I -----------------1 --------------------1 —
1 .4 0
— l
50
♦
♦
100
<
O
U
s
I
P
•H
O
U
O
CU
400
♦
♦
4
♦
+
too
4
4
4
4
4
4
.
X
.♦
X
X
1200
112
NALCO A CARRIER
♦
+
<■+
X
++4 4-4
4+
ISOTHERM FOR 8 8 5 3 4
L C 117-14 82 (1 1 3 7 )
S a m p le N o. 7 8 - 6 0 0 8A
400
+
300
+
2 00
+
X
100
.5
RELATIVE PRESSURE (P /P o )
1.0
GAS/CM CATALYST
4
113
83534 L C l17-1432
(1 1 3 7 )
4 -N o v -7 0
A D SO R PTIO N PORE VOLUME DATA
2 2 PO IN T PV AND 2 PO IN T SA
/*
BET SA =
12.00 P O R E
PDRE
D IA M ETER
A
X
1200
6 0 0
5 0 0
4 0 0
3 0 0
2 0 0
195
190
185
180
175
170
165
160
155
350
145
140
135
. 130
125
120
115
T-
no -
105 -
ICO -
950 -
9
8
8
7
7
6
6
5
5
. 4
5
0
5
0
5
0
5
0
5
—
—
—
40 -
603
500
400
300
203
195
I90
185
180
175
170
165
160
155
150
145
140
135
130
125
I20
.1 1 5
I IO
105
100
95
90
85
80
75
70
65
60
.55
50
45
40
35
'3 2 3 .2 3
VOLUME =
0 .7 1 8 3
SA T PV = 0 .7 6 4 3
4V /A =
8 8 .9
PORE
VOLUME
(C C /G )
C U M U LA TIV E
PORE VOLUME
(C C /G )
0 .0 3 6 5
0 .0 0 2 8
0. 0 3 4 1
0 .0 0 7 2
0 . 0 2 11
0 .0 0 2 6
0 .0 0 2 3
0 .0 0 3 2
0 .0 0 3 6
0.0 0 4 1
0 .0 0 4 6
0 .0 0 5 3
0 . OOoO
■ 0.0371
0 .0 3 3 2
0 .0 3 9 5
0.01.11
0 .0 1 2 9
0 .0154
0 .0 1 7 9
0 .0 2 1 3
0 .0 2 2 9
0 .0 2 4 6
0 .0 2 7 3
0 .0 3 1 3
0 .0 3 6 3
0 .0 4 0 9
0 .0 4 1 1
0 .0 4 2 0
0 .0 4 3 4
0 .0 4 3 1
0 .0 3 9 3
0 .0 3 3 2
0 .0 3 3 5
0 .0 2 9 9
0 .0 2 6 3
0 .0 2 2 3
0. 0331
0 .0 0 7
0 .0 0 9
0 .0 1 3
0 .0 2 1
0 .0 4 2
0 .0 4 4
0 .0 4 7
0 .0 5 0
0 .0 5 4
0 .0 5 3
0 .0 6 3
0 .0 6 3
0 .0 7 4
0 .0 3 1
0 .0 3 9
0 .0 9 9
O .H O
0 .1 2 3
0 .1 3 3
0 .1 5 6
0 .1 7 7
0 .2 0 3
0 .2 2 5
0 .2 5 3
0 .2 8 4
0 .3 2 0
0 .3 6 1 .
0 .4 0 2
0 .4 4 4
0 .4 3 3
0 .5 3 1
0 .5 7 0
0 .6 0 3
0 .6 3 7
0 .6 6 7
0 .6 9 3
0 .7 1 5
0 .7 1 8
.
■
M ED IA N
M EDIAN
VOL
PERCENT
(C C /G )
0 .9
0 .4
0 .6
1 .0
2 .9
0 .4
0 .4
0 .4
0 .5
0 .6
0 .6
0 .7
0 .8
1 .0
1 .1
1 .3
1 .5
1 .8
2 .1
2 .5
3 .0
3 .2
3 .4
3 .9
4 .4
5. I
5 .7
5 .7
5 .8
6 .0
6 .0
5 .5
4 .6
4 .7
4 .2
3 .7
3 .1
0 .4
0 .9
0 .4
0.6
1.0
2 .9
SURFACE
AREA
(S O M /G )
0 .2 7
0.20
0 .3 7
0 .8 4
3 .61
0 .5 2
0 .5 9
0.68
1 .7
2.8
5 .0
9 .4
1 4 .3
10.8
11 . 6
12,0
10. I
8.8
6.8
PD ON
PD ON
0 .7 9
0 .9 2
1 .0 8
1 .2 7
1 ,4 8
1 .8 0
2 .1 6
2 .5 8
3 .1 0
3 .7 5
4 .6 5
5 .6 3
6 .9 7
7 .7 9
8 .7 5
3 0 .3 5
1 2 . 22
1 4 .9 1
1 7 .6 3
1 8 .7 7
2 0 .3 7
2 2 .3 9
2 3 .7 9
2 3 .2 6
2 1 .2 7
2 3 .3 3
2 2 .7 7
22.12
2 0 .9 7
3 .2 6
SA =
PV =
7 6 .8
9 0 .2
C U M ULA TIV E
SURFACE
AREA
0 .3
0 .5
0 .8
$.7
lie
6 .4
7 . 1
7 .9
8.8
9 .9
11.1
12.6
1 4.4
I6 .6
1 9 .2
2 2 .3
2 6 . C
3 0 .7
3 6 .3
4 3 .3
5 1 .0
5 9 .6
7 0 . I
8 2 .4
9 7 .3
1 1 4 .9
13 3 .7
1 5 4 .1
1 7 6 .5
2 0 0 .3
2 2 3 .5
2 4 4 .8
2 6 8 . I
2 9 0 .9
3 1 3 .0
3 3 4 .0
3 3 7 .3
114
88534
L C I17-1432
( 1 137)
4 -N o v -7 3
ADSO RPTIO N
(D E L T A
0
PORE
V /
10
VOLUME DATA
D E L T A D ) * l OOO
20
1 — -------------------- 1-------------------------1------------------------ 1-------------------------1
20
40
60
100
120
140
160
180
200
PORE DIAMETER, ° A
80
115
88534
L C I1 7 -1482
(1 1 3 7 )
—
4 -N o v -78
D ESO R PTIO N PORE VOLUME DATA
21 PO IN T PV AND 2 PO IN T SA
BET SA S ' 3 2 3 .2 3
PORE VOLUME =
IZDO
PORE
D IA M ETER
A
1200
6 0 0
5 0 0
4 0 0
3 0 0
2 0 0
195
190
185
180
175
170
1 6 5
16 0
155
I5 0
145
140
135
830
125
120
115
HO
105
80 0
95
9
8
8
7
7
6
6
600
500
400
300
200
195
190
185
160
175
1.70
165
160
155
150
145
140
135
130
125
- 120
- 1 15
- H O
- 105
- 100
95
- 90
85
80
75
70
65
— 60
—
-
0
5
0
5
0
5
0 -
35 -
55
50
SA T PV = 0 . 7 6 4 8
4V /A =
91 .7
0 . 7414
PORE
VOLUME
( C C /G )
0 .0 0 2 6
0 .0 0 1 2
0 .0 0 1 2
0 .0 0 2 1
0 .0 0 3 0
0. 0002
0. 0003
0. 0003
0 . 0003
0. 0004
0. 0004
0. 0004\
0 .0 0 0 5
0. 0005
0. 0005
U. 0005
0. 0005
0 . OOOo
0 . OOOo
0. 0007
0. 0007
0 .0 0 0 9
0. 0009
0 .0 0 1 0
0 .0 0 1 0
0 .0 0 0 0
0 .0 0 6 2
0 . 11 9 2
0. 1343
0 .1 1 3 9
0 . I003
0 .0 3 6 3
0 .0 7 1 5
0 .0 6 1 6
0 .0 2 6 0
C U M U LA TIV E
PORE VOLUME
(C C /G )
0
0
‘0
0
0
0
0
.0
.0
.0
.0
.0
.0
.0
0 3
0 4
0 5
0 7
1 0
1 0
11
o.oii
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
.0 1 1
.0 1 2
.0 1 2
.0 1 2
.0 1 3
.0 1 3
.0 1 4
.0 1 4
.0 1 5
.0 1 5
.0 1 6
.0 1 7
.0 1 7
.0 1 8
.0 1 9
.0 2 0
.0 2 1
.0 2 1
.0 2 7
.1 4 7
.2 8 1
.3 9 5
.4 9 6
.5 8 2
.6 5 4
.7 1 5
.7 4 1
M EDIAN
M EDIAN
VOL
PERCENT
(C C /G )
0 .4
0 .2
0 .2
0 .3
0 .4
0 .0
0 .0
0 .0
0 .0
0. I
0 . I
0. I
0 . 1
0 . I
0. I
0. I
0 .1
0. I
0. I
0 . 1
0 . 1
0. I
0. I
0. I
0 . 1
0 .0
0 .8
16. I
1 8.2
1 5.4
13 .6
1 1 .6
9 .6
8 .3
3 .5
0
0
0
0
0
.
.
.
.
.
PD ON
PD ON
SURFACE
AREA
(S O M /G )
4
2
2
3
4
0 . 1
0 .2
0 .3
0 .4
0 .5
0 .8
3 4 .3
2 9 .0
2 1 .3
1 1 .8
0 .1 5
0 .0 9
0..11
0 .2 4
0 .5 0
0 .0 5
0 .0 5
0 .0 6
0 .0 7
0 .0 9
0 .0 9
0 . 10
O . . 11
0 .1 2
0 .1 3
0 . 14
0 . 15
0 . 17
0 .1 9
0.2.1
0 .2 4
0 .2 9
0 .31
0 .3 9
0 .3 7
0 .0 0
2 .6 7
5 4 .4 7
6 5 .3 7
5 8 .7 9
5 5 .5 9
5 1 .1 4
4 5 .7 6
4 2 .8 6
1 9 .7 9
SA =
PV =
7 3 .
7 6 .
C U M ULA TIV E
SURFACE
AREA
:
0 .2
0 .2
0 .3
0 .6
1.1
1 .1
1 .2
1 .2
1 .3
1 .4
1 .5
1 .6
1 .7
1 .8
2 .0
2. I
2 .3
2 .4
2 .6
2 .8
3. I
3 .4
3 .7
4 . I
4 .4
4 .4
7. I
6 1 .6
1 2 6 .9
1 8 5 .7
2 4 1 .3
2 9 2 .5
3 3 8 .2
3 8 1 .1
4 0 0 .9
116
8 8534
L C I1 7 -1 4 8 2
(1137)
4 —N o v —7 8
PORE
VOLUME DATA
ADSO RPTIO N
BET SA
= 3 2 3 .2
C U M U LA TIV E
SQ
D ESO RPTIO N
M /G
3 3 7 .3
SA
INTERCEPT =
0 .0 0 0 1 4 4
C U M ULA TIV E
SLOPE
SA
= 0 .0 1 3 3 1 3
4 0 0 .9
M EDIAN
PD ON SA
7 6 .8
M ED IA N
PD ON SA
7 3 .7
M EDIAN
PD ON
9 0 .2
M EDIAN
PD ON PV
7 6 .1
8 8 .9
AVERAGE
AVERAGE
P D C 4V /A )
s a t u r a t io n
D IA
PV
PV
PV
PV
PV
PV
PV
PV
PV
1203
1003
803 =
6 0 0 =
4 03
3 0 0
203
103
0 .7 6 5
PV
SATURATIO N
C C /G
.7
.7
.7
.7
.7
.6
.6
.4
1
1
1
1
0
9
7
3
8
4
2
1
4
7
6
4
3
3
6
7
3
o
o
3
P D C 4V /A )
D IA
PV
PV
PV
PV
PV
PV
PV
PV
1200
I 000
8 0 0 =
6 0 0
4 0 0
3 0 0
2 0 0
100
PV
C C /G
.7
.7
.7
.7
.7
.7
.7
.7
4
4
4
3
3
3
3
2
1
1
1
3
6
4
1
0
4
4
4
3
5
3
4
2
=
9 1 .7
=
0 .7 6 5
117
63534
L C M 7 - 14 8 2
(1137)
4 -N o v -7 8
SO RPTIO N
(D E L T A
0
PORE
V /
30
VOLUME DATA
DELTA D )*IO O O
6 0
1 ------------------ 1---------------------- 1---------------------- 1---------------------- 1
20
40
60
100
120
140
160
180
200
PORE DIAMETER,
°A
80
118
88534 LC 117-1482 (1 1 3 7 )
4 -H o v -7 u
<DELTA PORE VOL'JMEZDELTa LOO PORE DIAMEfER)
. 70
1 .4 0
2 .1 0
2 .0 0
3 .5 0
4 .2 0
—♦ ♦ ♦ ♦ X X X X X X X X X X
0
4 .9 0
25
50
♦
+
♦
4
4
4
4
200
X X X X X K H H H »<♦
40 0
♦
♦
.
.
600
1200
o*
PORE DIAIfETER
100
119
APPENDIX B
SAMPLE CALCULATION OF PORE VOLUMES
CATALYST PREPARED BY
THE IN C IP IE N T WETNESS METHOD
C ataly st d esired
MoO3 ( d e n s i t y
CoO
(density
WO3
(d en sity
4 . 6 9 g m /m l) :
6 . 4 5 g m /m l) :
7 . 1 6 gm /m l) :
%
4%
8%
1 0
B lank c a r r i e r a s r e c e i v e d
a fte r calcin atio n
volum e
nom inal P .V .*
A.
149.4
136.4
300
1 . 0.
gm
gm
ml
ml/gm
(I)
10% MoO3
The s o l u t i o n c o n c e n t r a t i o n ,
10% MoO3 i m p r e g n a t i o n i s
X, r e q u i r e d
____( P . V . ml) (X gm MoO3 / m l s o l . )
c a r r i e r gm
1 ^ Q0 X = 0 . 1
to g iv e
= q 10
X = 0 . 1 gm MoO3 / m l
To c o m p l e t e l y s o a k 300 ml o f c a r r i e r ,
200
ml o f s o l u t i o n .
we n e e d a b o u t
200 m l ( 0 . 1 gm /m l) = 20 gm MoO3
20 g m / 0 . 8 1 4
= 2 4 . 5 7 gm ammonium m o l y b d a t e
The r e s u l t i n g s o l u t i o n g i v e s a s p g r o f 1 . 0 8 6 .
Th e c a r r i e r w a s b r o u g h t i n t o c o n t a c t w i t h t h e
tio n and then d ra in e d the e x c e s s iv e s o l u t i o n .
Wt o f s o l u t i o n a b s o r b e d _ 1 4 4 . 1 gm _ Q
Wt o f i n i t i a l s o l u t i o n
2 1 7 . 2 gm
Wt o f MoO3 a b s o r b e d = 20 g m ( 0 . 6 6 3 )
13.269
= 0 . 0 9 7 o r 9.7%
136.4
Wt o f MoO3 a b s o r b e d
Wt o f c a r r i e r
E rror =
= 1 3 . 2 6 9 gm
x
100
= -3.0%
T o t a l w e ig h t a f t e r d r y i n g = 155.0
C arrier
136.4
MoO3
13.27
HgO
5 .33
gm
gm
gm
gm
solu­
121
' C a l c u l a t e d P . V . a f t e r MoO3 i m p r e g n a t i o n a n d d r y i n g
13.27
0
.
9 7
+ 5.32
* ------4 ^ f l ?
1 ------
4
Th e m e a s u r e d P . V .
=
0
. 9 1 ( m l / gm)
(2 )
is
1 4 4 . 1 gm
1 . 036 g m / m l =
1 3 6 . 4 gm
0 e 9 7
ml/gm*
(3)
This m easured v a lu e a g re e s w e ll w ith V a lu e (I ) .
The V a l u e (2) i s t o b e c h e c k e d b y t h e CoO i m p r e g n a ­
tio n .
B . 4% CoO
X = 0.04
X = 0 . 0 4 4 CoO gm /m l
200 m l ( 0 . 0 4 4 CoO gm /m l) = 8 . 7 9 1 gm CoO
8 . 7 9 1 gm C oO ( 1 / 0 . 3 2 5 5 ) ( 1 / 0 . 9 9 3 8 ) = 2 7 . 1 8 gm a s
c o b a lt n i t r a t e hexahydrate
Wt o f
Wt o f
Wt o f
5 .2 8
136.4
s o l u t i o n a b s o r b e d _ 1 2 9 . 5 gm = 0 . 6 0 1
2 1 5 . 6 gm
i n i t i a l so lu tio n
5 . 2 8 gm
CoO a b s o r b e d = 8 . 7 9 1 ( 0 . 6 0 1 )
gm
3.3%
x 100 = 3.87% CoO; e r r o r
gm
The m e a s u r e d P . V .
1 2 9 . 5 gm
1 . 0 7 8 gm/ml
1 3 6 . 4 gm
is
0 .8 8
(4)
m l/gm
C om paring t h i s w i t h V alu e ( 2 ) ,
error .
0
^
i
80 - 9 1
*
100
- -3.3%
1 6 5 . 7 gm
T o tal w eight a f t e r drying
1 3 6 . 4 gm
C arrier
1 3 . 2 7 gm
MoO 3
5 . 2 8 gm
CoO
1
0 . 7 5 gm
H20
1 3 .2 7 , 5 .2 8 , 10.75
4.69
6.45
0 .8 6
ml/gm
0 .9 7
136.4
122
Th e c a l c u l a t e d P »V. a f t e i r MoO3 a n d CoO i m p r e g n a t i ­
on and d r y i n g . i s 0 .8 6 ml/gm .
C
8% WO3
0 .86 X
0.0 8
X = 0 . 0 9 2 5 gm WO 3 / m l
I
150 m l s o l u t i o n ( 0 . 0 9 2 5 gm WO3 / m l ) = 1 3 . 8 8 gm WO3
1 3 . 8 8 g m ( I . 1 5 0 4 2 ) = 1 5 . 9 7 gm a s ammonium t u n g s t a t e
(meta)
The s p g r o f t h i s s o l u t i o n w a s 1 . 0 9 9 3 .
1 3 0 . 4 gm
0.7908
1 6 4 . 9 gm
W t o f WO3 a b s o r b e d = 1 3 . 8 8 g m ( 0 . 7 9 0 8 ) = 1 0 . 9 8 gm
Wt% i m p r e g n a t e d = 1 0 . 9 8 ( 1 0 0 ) / 1 3 6 . 4 = 8.05%
M e a s u re d P .V . = 0 . 8 7 ml/gm
e r r o r = -1.16%
The c a l c u l a t e d P . V . a f t e r MoO3 , CoO, a n d WO3
im p reg n atio n and d ry in g
1 3 .2 7 , 5 .2 8 , 10.98 , 13.07
4.69 ^ 6 .4 5
7.16
I
0 . 8 4 ml/gm
136.4
The f i n a l m e t a l c o m p o s i t i o n s
T o tal w eight
C arrier
MoO 3
CoO
WO 3
H20
a nd P.V,
165 I
136 4 ■
13, 27
5 28
1 0 , 98
- 0 . 83
N e g l e c t i n g th e u n a c c o u n te d l o s s (w ater)
1 3 .2 7 , 5 .2 8 , 10.98
4 .6 9 + 6 .4 5 + 7.16
0 .9 7
0 .93
136 1 4
a fte r
gm
gm
gm
gm
gm
gm
calc.
123
APPENDIX C
X st m d -86 d is t il l a t io n d a t a
ASTM DISTILLATION FOR THE KT SERIES CATALYSTS
—
KT-I
KT-2
—
—
364
386
390
392
396
410
620
620
360
380
384
390
396
407
(420)
-
-
T e m p e r a t u r e , uF
KT -4
KT-5
—
370
378
362
376
384
374
386
38 1
390
391
392
384
402
396
390
410
404
420
430
—
(656)
(660)
670
X
n
i
OJ
V ol.
KT-O
ml
243
IBP
382
5
386
10
390
15
- 39 4
20
2 5 ( 2 3 . 5) ( 4 3 6 )
30
—
35(34)
37(36)
40(39)
KT-6
—
368
374
384
390
396
408
442
KT-7
—
375
384
388
392
. 400
412
650
K T -8
—
362
378
392
388
394
402
648
—
—
—
—
KT-16
130
362
372
380
386
390
400
416
(590)
KT- 17
154
350
372
/378
384
390
400
410
610
124
—
658
V ol.
ml
KT-9
IB P
5
10
15
20
25
30
35
40(39)
K T - 10
120
364
374
380
386
390
39 8
415
590
356
374
380
384
390
396
420
590
K T-Il
14 2
368
378
382
390
396
410
620
-
T e m p e r a t u r e , °F
K T - 12 K T - 13 K T - 14
138
286
114
378
364
360
380
378
376
390
. 382
381
394
386
386
402
392
392
422
400
400
640
435
430
616
K T - 15
ITo
364
376
384
390
396
410
600
—
ASTM DISTILLATION FOR THE KT SERIES CATALYSTS-CONTINUED
V ol.
ml
IBP
5
10
15
20
25
30
33(32)
33.5
3 5 ( 3 4 . 5)
K T - 18
240
366
374
380
386
392
404
K T - 19
260
368
376
380
386
392
402
—
—
K T - 20
225
380
384
388
39 2
400
418
(438)
—
—
—
410
420
—
T em p e ra tu re f0F
KT-21 KT-22 KT-2 3
204
230
204
378
364
380
384
388
380
392
389
386
392
390
396
402
400 „
396
424
402
420
445
446
' 4 38
KT-24
244
380
386
388
392
398
418
436
K T- 25
230
376
382
386
392
400
418
450
• 590
KT-2 6
180
378
384
. 388
392
400
418
432
to
Ln
V ol.
ml
IBP
5
10
15
20
25
30
32
33.5
35
KT-2 7
238
378
386
390
39 2
402
422
—
580
KT-IR
170
364
3 85
390
394
396
409
—
_
-
KT-3R
16 8
360
375
380
393
39 1
444
446
—
-
“
T e m p e r a t u r e , "Op
KT-7R K T - 9 R KT -1 4R KT- 19R K T - 2 IR KT-25R
2 32
250
230
22 8
210
222
36 7
36 5
375
377
364
373
381
375
380
379
373
385
384
385
392
381
382
389
388
386
391
388
395
39 2
394
394
395
406
39 1
401
402
442
400
400
412
400
—
442
456
43 8
610
—
■—
620
. 650
ASTM DISTILLATION FOR THE KT SERIES CATALYSTS-CONTINUED
V ol.
ml
IBP
5
10
15
20
25
30
33.5
37.5
T e m p e r a t u r e , °F
K T - 2 7R
237
377
386 .
390
394
400
416
4 32
620
to
ON
127
APPENDIX D
MODELING OF KATALCO CARRIER
128
MODELING OF KATALCO CARRIER
I.
P h y sic a l C h a r a c te r is tic s o f K atalco C a rrie r
C h e m i c a l c o m p o n e n t s : 98% AlgO g
2% S i0 2
True d e n s i t y ;
3 . 9 g m /m l
S o l i d volum e:
0 . 2 5 6 4 1 ml/gm
P o r e volum e:
0 .9 3 3 m l/gm
P o re v o l . / s o l i d v o l . : 3.64
Surface area:
207 m2/ g m
2.
BaSe G r a n u l e :
B a s i s : I gm
(I)
S p h e re w i t h sm ooth s u r f a c e
0 . 2 5 6 4 1 (m l /g m )
10**6
(m3/ CItl3)
m 3)
4 Tt ( § > * * 3 ( ^ 1 1 )
b a l l 10**6 'em '
(2)
4 7r ( 5 ) * ‘ 2 ( B| g ) T5^ ( ^ z ) N ( b2 aa li li sB, )
S o lv in g th e above 2 e q u a tio n s
D = 74°A
N = 1 . 193*10**18
3 » S S anning E l s g t r o n
N
(b a lls
' gm
!2 0 1 < ls>
sim ultaneously
Ph© t © m i @ r © g r a p h
I mm ^ 200©A
B 3,64 (=por@ ygl,/solid vol,)
M = 3 ,1
(n u n ^ e z o f b i l l s
o o o r d in its)
'
in
3-D
MONTANA STATE UNIVERSITY LIBRARIES
stks 0378.K56135@Theses
Catalysts for upgrading solvent refined
3
1762 00169312 4
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