Table Of Contentt i i @ !iilHiHi i iHii iH H iiiiii!il
ehT dn2 International Symposium no Hydrotreatment dna Hydrocracking of Oil ,snoitcarF
which si also eht ht7 ni eht seires of naeporuE spohskroW no Hydrotreatment, took place ni
Antwerpen, Belgium from November 41 ot .71
The Symposium emphasized how oil refining faces increasingly severe environmental
regulations. These dna the increasing application of heavier crudes containing more ,-S
N- dna metal components call for more efficient hydrotreatment dna hydrocracking
.sessecorp tI si clear from the keynote lectures, the oral contributions dna the posters of
this meeting that adapting the operating conditions will not suffice. Adequate catalysts
need to eb developed, with different composition dna structure. Surface science
techniques and molecular modeling era now well established tools for such a
development. They should eb of help ni widely different aspects, like the role of
precursors ni the preparation or the modifications undergone by the catalyst under
reaction conditions.
The improvement of hydrotreatment dna hydrocracking also needs accurate modeling
of the chemical reactor. This requires more representative hydrodynamics and kinetic
models whose validity extend to the very low S-and N-contents. These areas should eb
vigorously developed.
We look back ta a successful symposium with contributions from all over the world,
reflecting the state of the art ni industrial practice, ni industrial research centers and ni
academia. teL these Proceedings dissiminate the information presented ta the
Symposium also to those who were not able to attend.
B.Delmon, Universit~ Catholique de Louvain,Belgium
G.F.Froment, Texas A & M University, ASU
P.Grange ,Universit~ Catholique de Louvain,Belgium
The 2nd International Symposium "Hydrotreatment dna Hydrocracking of Oil Fractions"
saw organized by : The Technological Institute associated with the Royal Flemish
Society of Engineers IT( - K VIV).
The Technological Institute saw founded ni 1940 with the aim of disseminating
information on scientific dna technological development by means of seminars,
lectures, courses dna conferences.
Address :Technological Institute vzw
Desguinlei 214, B- 2018 Antwerpen Ir
tel: +3232160996
fax: +3232160689
e-mail : [email protected]
Hydrotreatment dna Hydrocracking of liO snoitcarF
.B Delmon, G.F. Froment dna .P Grange )srotidE(
(cid:14)9 1999 Elsevier Science B.V. All rights .devreser
Hydrocracking in the Year 2000: A strong interaction between
Technology Development and Market Requirements
(J.K. Minderhoud, J.A.R. van Veen and A.P. Hagan)
Summary
New developments in hydrocracking technology are increasingly guided by the prevailing market
conditions and requirements. This necessitates a very good knowledge of local circumstances in
the refinery as well as all integrated approach of the various disciplines involved such as catalyst,
process, reactor and engineering technology. Some recent developments, illustrated by a case
study, are discussed.
noitcudortnI
Starting in the early 1960's, hydrocracking has become one of the major conversion processes in
the refinery. It usually converts a rather heavy, low quality feedstock into lighter, highly valuable
transportation fuels, contributing significantly to the overall profitability of the refinery.
Due to the nature of the process, hydrocracking is predominantly suited to producing middle
distillates with excellent product qualities. Jet and diesel fractions can be obtained with very low
sulphur contents (often below 20 ppm) and very good combustion properties (kerosine smoke
points above 25 rain and diesel cetane numbers above 55). The obvious reason for it is the
relatively high hydrogen pressure used, typically above 100 bar, which results in high removal
rates of hetero-atoms (sulphur, nitrogen) contained in the feedstock and deep saturation of
aromatic compounds.
Other typical characteristics of hydrocracking are its flexibility in varying product slate, depending
to a large extent on the type of catalysts used, and its potential to produce very good quality
feedstock for lube base oil manufacturing, ethylene crackers and fluid catalytic crackers.
Full conversion hydrocrackers usually contain at least two catalysts: a pretreatment and a cracking
catalyst. Both catalysts are bifunctional: they contain a hydrogenation function and an acidic
function (see Table .)1 Pretreatlnent catalysts combine a strong hydrogenation function with a
rather weak acidic function, whereas for cracking catalysts both functions should be well balanced.
By the end of this century, the world hydrocracking capacity will be about 200 million tons per
annum. Projections were that more than 225 million tons per annum of processing capacity would
be available by the year 2002, with the highest growth rate expected to be in the Asia-Pacific zone,
but the recent economic downturns have resulted in a slower increase in capacity. Currently, it is
more realistic to expect a 3-5% growth in the next five years.
The main driving forces for the expansion of the hydrocracking process are :
(cid:12)9 a steady, but continuing growth in middle distillates consumption in all parts of the world.
(cid:12)9 introduction of ever more stringent automotive fuel specifications (sulphur, aromatics, 95%vol
recovery temperature) as prescribed in the USA and Europe.
(cid:12)9 processing synergy in combining catalytic cracking with hydrocracking, two large scale
processes
(cid:12)9 increasing development of mild hydrocracking technology 1,2
Technology Development
Over the past forty years, a large number of different process and reactor configurations have been
developed to carry out the hydrocracking process (Figure )1 3-6. Key differences have been:
(cid:12)9 two stage versus single stage / series flow operation
(cid:12)9 once through versus recycle mode of operation
(cid:12)9 a common versus a separate recycle gas system for first and second stage
(cid:12)9 world-scale single reactors containing multiple catalysts versus two or more reactors containing
individual catalysts
To comply with the demand for new hydrocracking capacity, it is evident that building new
hydrocrackers will continue in the years to come, despite the likely incremental increase of
capacity in existing hydrocracking units.
In an economic climate of low refining margins and emphasis on high returns on investment, there
is a very strong incentive to design and construct hydrocrackers with minimum capital investment.
It often means "silnple" units, i.e. single reactors, operating in once-through mode and at low
pressures.This is the major reason that the mild hydrocracking process has received increased
attention in the last decade.
On the other hand, the call for ultra low sulphur and, in particular, very low aromatics levels in the
products cannot easily be satisfied by applying a low hydrogen partial pressure process. Moreover,
catalyst activities in mild hydrocracking are reduced as well, leading to higher catalyst volumes,
i.e. bigger reactors, to achieve the same feed conversion level. These effects are illustrated in
Table 2.
In selecting the appropriate hydrocracking process, refiners will strive for the most cost-efficient
way to achieve their objectives. For hydrocracking process licensors this implies that they need to
strike the right balance between expensive, complex, high pressure processes offering much
flexibility, products of superior quality, and cheap, simple, low pressure designs with more
restricted deliverables. An example of this is the single-stage, single-reactor, stacked bed line-up
with optional liquid recycle (Figure 2). With advanced heat integration, a four separator reactor
effluent system and less fractionation equipment, this design offers clear advantages over the
oldest, conventional two stage processes.
In view of strong pressures to increase refinery margins, an increasing interest in optimising and
revamping existing hydrocrackers has been shown in the last few years and is expected to continue
in the next decade. There are numerous topics that can be addressed to achieve this goal (Table 3).
In this article, we will highlight (i) the use of new and improved catalysts, (ii) proper arrangement
of different catalysts in stacked bed configuration and (iii) new developments in reactor internals
and fouling abatement.
Finally, to emphasise the importance of close co-operation between refiner and process developer,
we will present a case study in which various process aspects played a very important role in
revamp and catalyst selection of the hydrocracker.
New and improved hydrocracking catalysts
In present day industrial research, effectiveness and efficiency are key, and catalyst research and
development is no exception to this. Target setting and the way R&D programmes are executed
have received much attention in recent years.
It is crucial that the targets for a new catalyst, in terms of improved performance and/or reduced
costs, are valued by the final user. To achieve this, targets are increasingly defined in close co-
operation with refiners in order to include their specific requirements. For locations where a long
term relationship has been developed, there is even a trend of tailor made catalysts.
As a result, catalyst companies and process licensors continue to introduce new hydrocracking
catalysts to the market (7-11). Here, we will discuss some new catalysts that were recently
developed by Criterion and ZI, taking into account refiners' wishes.
DN-190: A high activity pretreatment catalyst
A considerable number of hydrocrackers are severely constrained by an inability to meet the
required nitrogen slip to the cracking catalyst, often aggravated by the processing of heavier
feedstocks or feedstocks containing nitrogen species difficult to hydrodenitrogenate. Without
better pretreat catalyst, these hydrocrackers suffer from shorter cycle lengths.
To develop new pretreatment catalysts three widely different development concepts have been in
use: (i) high dispersion, (ii) controlled acidity, and (iii) optimised pore structure. The overall
concept is depicted in Figure 3.
DN-190 was developed on the basis of a high dispersion of the catalytically active phase. The
concept behind DN-190 is to boost the HDN activity by increasing the number of active sites by
enhancing the catalyst surface area on a reactor volume basis. This volumetric catalyst surface
area can be favoured by:
(cid:12)9 enhancing the CBD by optimising the shape and reducing the size of the extrudates. The latter
favours mass transport, too. To avoid excessive pressure drop over the reactor, there is a
minimum size and shape to ensure that the void fraction is at least 40%.
(cid:12)9 adjusting the textural properties of the carrier, according to:
SA = F * PV /MPoD (1)
where F is a numerical constant which depends on the shape of the pores (4 in the case of
cylindrical pores), PV the pore volume, and MPoD the median pore diameter.
DN-190 is based on Century technology, a process in which a nano-crystalline phase of alumina
TM
is synthesised in-situ on the gamlna-alumina support to generate slit-shaped pores. The nano-
crystalline phase suppresses stacking of molybdenum sulphide in the working catalyst. The
existence of single layers leads to a very high dispersion of the active phase 12.
The result is a catalyst displaying a very high volumetric surface area. Table 4 clearly confirms
that, as a consequence, the RVA of DN-190 is substantially higher than that of a conventional
pretreat catalyst such as C-424.
".326-Z A high active, high middle distillate selective zeolitic catalyst
It is well known that for cracking catalysts there is a kind of trade-off between high activity, high
naphtha selectivity and low activity, high middle distillate selectivity 7. The objective of new
cracking catalysts is often to improve on activity whilst maintaining selectivity or vice versa. This
can be achieved by altering the acidic function and/or the hydrogenation function. Zeolite
modifications for instance are numerous: dealulnination, insertion of silica, re-insertion of
alumina, morphology changes etc 13. Despite the fact that an overwhelming number of
treatments and modification routes have been explored already, there still appears to be room for
new and successful zeolite types. This led, as an example, to the development of a dealuminated Y
zeolite, which in combination with an amorphous silica-alumina (ASA) and a hydrogenation
function, finally resulted in the highly middle distillate selective catalyst Z-603 14.
Recently, by careful modification of both functions, Z-623 was developed as a result of efforts,
stimulated by the market, to obtain a catalyst with higher activity than Z-603 without
comprolnising on selectivity. Figure 4 demonstrates that the Z-623 performance is indeed better
than could normally be expected.
503.'A high diesel selective single stage catalyst
Catalysts using amorphous silica-alumina as acidic function are very well suited to maximising
diesel production. Furthermore, they have the ability for use in pretreating as well, because of the
relatively modest acidity. Consequently, this type of catalysts can be employed as the single
catalyst in a hydrocracker unit.
Improvements in ASA catalysts can be obtained by variations in e.g., composition, synthesis
conditions of the ASA, metal emplacement methods, post-treatments. Figure 5 shows how the
efficiency of the NiW hydrogenation function can heavily depend on the metal emplacelnent route
selected, which ultimately dictates the effectivity of the ASA catalyst in performance terms.
Increasing acidity of the ASA si in the first instance beneficial, but finally the performance
becomes less attractive due to an imbalance between acidity and maximum achievable
hydrogenation power.
In the development of 503, both the acidity of the ASA and the hydrogenation activity of the NiW
function were improved. The results are shown in Table 5, where the new 503 catalyst is compared
with the previous generation DW 800 catalyst: clearly, activity, selectivity and product properties
have been improved. The opportunity was also taken to manufacture a Ni/Mo analogue on this
new ASA support, designated 505. Due to its better HDN activity, 505 is best used for MHC duty
where any ilnprovement in HDN is beneficial in enhancing the cracking conversion. For high
pressure units, to make high quality middle distillates, 503 will be the higher performance catalyst,
though.
Catalyst stacked bed arrangements
To maximise the overall capacity and conversion capability of the hydrocracking catalyst system
or to minimise the overall catalyst volume, it si of paramount importance to optilnise the ratio of
pretreat over cracking catalyst.
A crucial factor in determining the optimum ratio is knowing the hydrocracking activity (reaction
rate constant for cracking) of the cracking catalyst as a function of pretreatment severity. Since
cracking catalysts are acidic and variations in pretreatment severity result in effluents containing
different amounts of organic nitrogen compounds, adsorption of basic nitrogen species can have a
dominant effect on the apparent cracking rate constant 15. This is further illustrated by
laboratory tests in which feedstocks with nitrogen contents varying from 20 to 280 ppmwt were
processed over a NiW/Dealuminated Zeolite catalyst. Figure 6 shows the following:
* there is a significant reduction of the apparent reaction rate constant for cracking with
increasing nitrogen content of the feed
, the penalty of increasing nitrogen content is (slowly) decreasing with increasing operating
temperatures
These effects are best understood by inhibition effects of adsorbed nitrogen which are described by
Langmuir-Hinshelwood rate expressions. Apart from suppressing the cracking reactions, the
nitrogen compounds also cause self-inhibition of the hydrodenitrogenation (HDN) reaction, as
shown in Figure .7 This effect in itself further retards the cracking reactions. It is also well known
that ammonia, organic sulphur compounds, hydrogen sulphide and (poly) aromatics display
inhibitive effects, but their adsorption constants are at least an order of magnitude lower than those
of organic nitrogen compounds 16.
The practical implications of the poisoning effects of nitrogen species on cracking catalysts are
very dependent on the operating conditions of the hydrocracker. Feedstock type, hydrogen partial
pressure and temperature window between start-of-run and end-of-run of the hydrocracker cycle
largely dictate what nitrogen content is acceptable in the feed entering the cracking catalyst. In two
stage units, due to the absence or very low levels of ammonia, operating temperatures in the
second stage are usually well below 400~ resulting in rather high nitrogen sensitivities. In single
stage (series flow) hydrocrackers, however, operating temperatures are generally (much) higher
and, hence, the cracking catalysts in those units can tolerate higher nitrogen slips from the
pretreating stage.
For the same reason, in two stage units, it is useful to consider applying stacked beds in the first
stage where temperatures are often higher than in the second stage. This may lead to higher
cracking conversion levels in the first stage, off-loading the duty in the second stage, as illustrated
ni Table 6. Overall, this will result in longer cycle run lengths.
Optimising the balance between the various catalysts in a hydrocracker calls for a good description
of all kinetic parameters. Process models are therefore an indispensable tool in improving the
operation of existing hydrocrackers and designing new ones. The type of models that have been
developed, vary from very rigorous and fundamental to simple and correlative 17-22.
Commercial models try to combine simplicity and user-friendliness with accuracy and
thoroughness.
Reactor internals
To achieve maximum utilisation of the catalyst inventory in hydrocrackers, it is essential to obtain
both an even radial distribution of liquid and gas across the catalyst beds and excellent interbed
quench performance 23.
Recently, various studies have been undertaken to better understand and improve the performance
of liquid distribution trays 24. Patel et. al. described the development of the so-called Vapor-Lift
Distribution Tray, which is claimed to have a much more stable, low tilt-sensitivity operation over
a wide range of vapour/liquid ratios than classic bubble cap trays.
The Shell developed, so-called High Dispersion (HD) trays were found to display a much better
liquid distribution uniformity than more conventional trays, as shown in Figure 8 25. Main
reason for the better performance of the HD tray is that contrary to conventional trays, gas and
liquid are passing together through the nozzle, which causes an acceleration of the liquid and an
intimate mixing of liquid and gas. Because of the excellent distribution at the top of the bed there
is no need for a layer of distributive packing above the catalyst bed, resulting in increased catalyst
volume per reactor.
Due to the exothernlicity of the hydrocracking reactions, it is necessary to apply interbed
quenching in order to achieve a safe and controlled operation and to optimise the axial reactor
temperature profile. Interbed internals, used to reach this goal, allow injection of a cold gas or
liquid medium and also need to provide adequate mixing of reactant liquids and gases toensure a
homogeneous, radial temperature profile at the top of the catalyst bed.
Many different interbed quench devices have been developed. One of the important aspects in new
designs, apart from providing good mixing, is the aim of minimising the height of the internal to
maximise the amount of catalyst to be loaded in the reactor. The Ultra Flat Quench (UFQ) internal
with a height of only 1 meter, described by Ouwerkerk et. al., elegantly meets this requirement
(Figure 9) 25. In cold-flow testing with an imposed 30~ temperature difference above the
internal, they reported a maximum radial delta T of only some 4~ below the internal in case of
the UFQ instead of 16~ for a conventional internal (Figure 10). The high performance of the
UFQ has been well demonstrated in several commercial operations.
Fouling abatement
An important aspect to extend the run length of a commercial hydrocracker is to prevent reactor
fouling since it will lead to increase of pressure drop and finally to a premature shut down. Fouling
is often caused by inert solid particles entrained in the feed that deposit on and between the
catalyst particles (salts, iron scale) or chemical substances that react and deposit on and between
the catalyst particles.
One of the elements to combat fouling is to apply graded layer loading of inert materials and
catalysts which is based on the concept of deep bed filtration 26. It entails loading higher
voidage, larger particle size materials in the top layer, followed by layers of gradually smaller
sized materials and finally the hydrocracking catalyst(s).
A potential disadvantage of using bed grading is the loss of reactor volume for loading the actual
hydrocracking catalyst since (a part of) the grading material can be inert (Raschig rings are often
used). To mitigate these effects, materials displaying some catalytic activity are now being used. It
should be realised, however, that, due to the larger particle size of the grading materials, diffusion
limitational effects are more pronounced, resulting in lower effective reaction rate constants.
Clearly, the key issue in optimising bed grading is finding the right balance between fouling
prevention and preserving sufficient overall catalyst activity.
Case Study
A further illustration of the many factors that count in improving hydrocracker operations is an
example of an existing partial conversion hydrocracker for which plans were made to increase feed
throughput by some 10%, to include some 10% more aromatic, higher nitrogen feed in the total
feed diet and to abandon using a halide agent as activity booster for the catalysts.
Moreover, the refinery had stated a number of premises that needed to be met: maintaining the
same runlength as before, limiting the extra amount of naphtha and lighter products to maximum
20% more, and producing unconverted oil (called hydrowax) with preferably the same quality as
before. It was realised that some revamping of the hydrocracker was inevitable, but the objectives
had to be met at minimum capital investment.
From a technology point of view, the requirements had the following impact:
(cid:12)9 In order to obtain the same quality hydrowax, feed conversion needed to stay at least at the
same level. This is illustrated in Figure 11 showing that BMCI, which is a measure for
hydrowax quality, deteriorates, for a given catalyst system and feed composition, with
decreasing conversion (note that a higher BMCI value corresponds with a lower quality).
(cid:12)9 The higher throughput, the more difficult feed to process and the abandoning of catalyst
activity booster, combined with the need to achieve the same cycle life at equivalent conversion
level, called for almost a 100% increase of the intrinsic activity of the catalyst system.
(cid:12)9 Higher catalyst activities could, in principle, be attained by switching to new, more active
pretreat and cracking catalysts. For the cracking catalyst, this meant a system containing a
zeolite in higher amounts and/or having a higher intrinsic activity. But, as shown in Figure 12,
such catalysts would definitely result in higher naphtha yields. The more so, since the extra
naphtha make due to the higher amount of feed processed had to be absorbed as well.
Moreover, with those catalysts, the hydrowax quality would also deteriorate (Figure 11).
In the selection of catalysts, optimising the ratio between pretreat and cracking catalyst played an
important role. Not only activity, yield and product quality aspects needed to be considered, but
also catalyst activity decline rates and individual bed quench capacities had to be taken into
account. Figure 31 indicates that the theoretical optimum is not always fitting with a discrete
number of catalyst beds, further complicating the final choice to be made.
On top of these process aspects, additional hardware related equipment became critical too. The
capacity of the fresh and recycle gas compressor, the feed furnace, the separators, the fractionator
work-up section and the heat exchangers had to be checked in order to review whether they could
cope with the more severe duty.
The revamp project that started was a multi-disciplinary approach with close co-operation between
refiner, technology provider and catalyst vendor. Feedback of detailed data from commercial
operation for tuning of the process model and exploiting the new options for the hydrocracker to
the limit of equipment capabilities, was a key step in successful execution of the project.
The final option selected centred on a revamp of the fractionator allowing more naphtha draw-off,
a choice for the more active DN-190 as pretreat catalyst and the more active Z-623 as cracking
catalyst, and replacement of one bed of pretreat byr cracking catalyst.
Conclusions
Since its origin in the 1960's, hydrocracking has become an important process in the refinery.
Over the years many improvements in hydrocracking technology have been implemented, both
from a process and a catalyst point of view. Although hydrocracking is a mature process, still new
developments come to the fore, stimulated by a steady, further growth of the market and
environmental pressures on product qualities.
More than ever, hydrocracking technology developments in industry are nowadays based on a
thorough knowledge of existing commercial operations. New catalysts are being developed not
just because they have a superior performance over older ones, but because they maximise
profitability in the refinery. Process developments are geared towards optilnising integration with
other refinery processes via processing lower quality feedstocks from catalytic crackers, thermal
crackers, residue conversion units etc., and producing high quality transportation fuels and
feedstocks for ethylene crackers, catalytic crackers and lube base oil plants.
The number of new hydrocrackers to be build in the next few years will be limited, caused by
depressed refinery margins. A major part of the market will lie in optimising existing
hydrocrackers. This requires attention to optimising catalyst packages, prevention of fouling,
maximum utilisation of catalyst reactor volume, careful feedstock selection, but also revamping of
equipment such as furnaces, gas compressors, separators, heat exchangers and fractionators.
Accurate and detailed process models are a prerequisite to identify tailor-made solutions. Such
models need to be based on both kinetic information from R&D experiments and data from
commercial operation. Close co-operation between refiner, technology provider and catalyst
vendor is a key factor in combining technology developments with market requirements.
Acknowledgement
The authors wish to express their thanks to C.E.D. Ouwerkerk, M.C. Zonnevylle and J.R.
Newsome from Shell Global Solutions, SIOP BV, Amsterdam, and W.H.J. Stork from Shell
International Chemicals BV, Amsterdam, the Netherlands for their contributions to this paper.
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