A History of Fischer-Tropsch Wax Upgrading at BP—from Catalyst Screening Studies to Full Scale Demonstration in Alaska

Available online a t www.sciencedirect.com

Journal of Natural Gas Chemistry 15(2006)1-10

SCIENCE PRESS

www.elsevier.comllocateijngc

Review

A History of Fischer-Tropsch Wax Upgrading at BPfrom Catalyst Screening Studies t o f i l l Scale Demonstration in Alaska John P. Collins', Joep J. H. M. Font

Freide2*,

Barry Nay2

1. B P Exploration America Inc., 501 Westlake Park Blvd., Houston, T X 77079, United States of America

2. B P Exploration Operating Company, Chertsey Road, Sunbury on Thames, Middlesex, T W 1 6 7 L N , United Kingdom

[ Manuscript received November 23, 2005; revised January 24, 2006 ]

Abstract: Conversion of Fischer-Tropsch wax into high quality synthetic crude or finished transportation fuels such as premium diesel has been studied over the past 15 years within BP. Catalyst screening and selection was carried out in dedicated micro-reactors and pilot plants, whose designs are critical to the performance selection. Variation in catalyst composition and defining the gas to oil feed ratios with the operating temperature are a few of the parameters studied. Product selection and maximizing diesel yield combined with stability (catalyst life) were the ultimate drivers. The selected catalyst was then tested under commercial conditions in a dedicated 300 barrel per day demonstration plant. The products were also tested in engines to assess their combustion characteristics. Key words: Fischer-Tropsch; wax upgrading; hydrocracking; synthetic crude; diesel; naphtha; engine; emission

1. Introduction There is growing interest in Gas to Liquids (GTL) conversion as a tool for monetizing natural gas and production of premium quality products free of nitrogen, sulfur, aromatics, and metals. Virtually unlimited markets already exist for GTL diesel engines since it can be used as a transportation fuel directly or blended to improve the performance of conventional diesels. Other ultra high purity GTL products include naphtha, jet fuel, kerosene, and lubricant based oils. In locations where processing into end products is not the preferred option, synthetic GTL crude may be transported to a refinery by pipeline, ship, or truck, alone or blended with crude oil. An opportunity is also provided to eliminate green house gas emissions from the flaring of associated gas while producing ultra clean products. GTL conversion involves an integrated process * Corresponding author. E-mail: [email protected]

consisting of the following three steps: 0 The conversion of natural gas into synthesis gas (Hz/CO) by steam or oxygen/air based reforming. 0 The conversion of synthesis gas into wax by the Fischer-Tropsch (FT) process. Wax upgrading to synthetic crude by hydrocracking followed by separation into the end products. The objective of the reforming and FT conversion steps is to produce a paraffin wax of suitable composition at high process selectivity while the final hydrocracking step is crucial for completing the conversion to synthetic crude or finished products. Catalyst performance is a key factor in both the FT and hydrocracking operations. The story of BP's FT catalyst development efforts was reported in a previous publication [ 11. Since no commercial catalysts were available, the FT catalyst program included the development of both a catalyst formulation and a commercial catalyst manufacturing process. Here we chronicle the search

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John P. Collins et a]./ Journal of Natural Gas Chemistry Vol. 15 No. 1 2006

for a wax upgrading hydrocracking catalyst from initial screening studies conducted in the early 1990 s to BP's GTL demonstration facility in Nikiski, Alaska. A future article will report on the progress of ongoing catalyst assessments conducted in collaboration with Albemarle Catalysts Company BV. 2. Catalyst selection criteria

The search for a hydrocracking catalyst began in early 1990 with the objective of producing a pumpable synthetic crude or middle distillates by hydrocracking FT wax. The initial experiments were aimed at demonstrating the feasibility of upgrading wax feed stocks by hydrocracking, screening studies to optimize the choice of catalyst and an assessment of the product quality. A process design and economic study was performed to establish performance targets that could be used to screen different catalysts under specified reactor conditions. Performance targets were established for wax conversion, product selectivities, and product quality indicators such as density, cetane number, and cold flow properties. Product selectivities were defined in terms of the relative conversion of 360 "C+ boiling point components in the wax to diesel/gas oil (250-360 "C), kerosene (150250 "C), naphtha (C5-150 "C), and gas (c1-C~)fractions. An overall objective was to maximize the selectivity to the gas oil (diesel) and kerosene middle distillate fractions. Catalyst activity and stability were other important considerations as the performance targets were specified for commercially relevant gas and liquid hourly space velocities (GHSV and LHSV) over a prolonged (2 year) period of operation. Blending studies were also conducted to confirm that synthetic crude produced from the selected catalyst was suitable for pipeline transport. The performance of the GTL diesel produced in the study and blends of the GTL diesel with conventional diesels was also investigated in engine emission tests. Due to the large number of conventional hydrocracking catalysts already in use, there was a pool of commercial catalysts to chose from for upgrading the FT wax. The FT wax differs from conventional hydrocracking feedstocks in that it contains no sulfur, nitrogen, or metals and the only hydrocarbon types present are n-paraffins. This influences the choice of catalyst for the wax upgrading. Catalyst candidates for the screening studies were selected with the assistance of internal hydrocracking expertise on petroleum waxes supplemented by a literature review. It

was concluded that for the case of FT wax upgrading, a relatively mild hydrocracking regime applies with lower pressures than those used in conventional hydrocracking. The choice of conditions for maximizing selectivity to middle distillates, particularly diesel, depends on the activity of the catalyst system and the desired conversion per pass. Careful consideration of the reaction mechanism for paraffin hydrocracking was needed to guide the selection of catalyst candidates. Mechanistic studies on normal alkane hydrocracking helped to identify desirable features of the catalyst functional balance [2-51. Since the initial steps are dehydrogenation followed by acid rearrangements and cracking of the intermediates, strongly dehydrogenating metals and strong acids can induce high conversions at low pressures. However, these two functions must be properly balanced with the hydrogenation function of the metal which saturates olefinic intermediates before they can crack further, thereby reducing over-cracking and optimizing selectivity for higher boiling range products such as diesel. Catalysts with a strong dehydrogenation function (noble metals) require relatively mild acidity to optimize selectivity to middle distillates while catalysts with a weaker dehydrogenation function can tolerate a stronger acid. The evaluated catalysts were chosen to provide strong dehydrogenation/hydrogenation functions in combination with a mild acid function. Catalyst systems which show this functional balance had been reported as the most suitable for maximizing selectivity to middle distillates [4]. Sufficient hydroisomerization activity was also a consideration since the cold flow properties of iso-paraffins are better than normal-paraffins. The catalysts studied included proprietary BP catalysts and other commercially available catalysts that were already in use or under assessment for use in conventional BP refining operations. Both base and noble metal catalysts were selected for evaluation.

3. Experimental The catalyst screening studies, started in 1990, were carried out in 20 ml micro-reactors using ground up catalyst particles. Several micro-catalytic catalyst evaluation (MCE) units were utilized. Model wax feeds consisting of a commercially available paraffin wax blend were studied for the most part. The catalyst selected in the MCE studies was then tested with the BP FT wax, In 1991, a 100 ml pilot plant (Pilot 1)

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Journal of Natural Gas Chemistr.y Vol. 15 No. 1 2006

was modified for whole particle evaluations of the top candidate from the MCE studies to confirm the catalyst's performance. A second 200 ml pilot plant (Pilot 2) was commissioned in 1992 for further evaluations of whole catalyst particles to overcome performance and operability problems encountered with the Pilot 1

unit. Reactor and catalyst dimensions for the microreactors and two pilot reactors are compared in Table 1. All units were operated in a single pass, down-flow trickle bed reactor configuration on once-through hydrogen in the screening studies.

T a b l e 1. C o m p a r i s o n of screening reactor dimensions Catalyst volume (cm3)

Diluent

Bed length (llcm)

Effective diameter (dlcm)

Ild

Catalyst particle diameter (Dp/mm)

1/D,

Furnace

420-150

single zone

Micro reactors

20

catalyst 1:2

21

1.10

19

0.5-1.4

Pilot reactor 1

100

1:2

30

2.06

15

3.175

95

3 zone

Pilot reactor 2

200

1:2

117

1.48

79

3.175

368

6 zone

The composition of the liquid product was determined using simulated distillation (SIMDIS) while vent and off-gases were analyzed by gas chromatography using instruments fitted with both thermal conductivity and flame ionization detectors. For product quality evaluations, several samples of liquid product collected over a period of time were combined and fractionated on a 14 plate distillation system. Combustion and cold flow properties of the fractionated products were then determined using standard methods.

draining of the two separators was required to allow safe collection of products and to prevent depressurization of the plant. The hydrocarbon product from the hot high pressure separator was collected by first transferring into a low pressure collection vessel followed by draining through a metering valve. Product from the cooled high pressure separator was collected by sequential manipulation of two valves which discharged aliquot amounts of the product into a collection container. The gaseous effluent composition was determined by analysis of the vent gas and the volume was measured by a gas meter.

3.1. Micro-reactor units 3.2. Pilot units The test apparatus consisted of a feed section, reactor module and product separation/product collection sections. The reactor module was a metal tube surrounded by an aluminum bronze block in a single zone furnace. An axial thermocouple was used for the internal temperature measurements. Temperature was controlled by a thermocouple placed between the reactor and aluminum bronze block. The catalyst charge consisted of 20 ml of crushed catalyst diluted with 10 ml of acid washed sand (0.5 mm diameter). Central location of the charge in the metal tube was accomplished by sandwiching it between inert ball sock beads separated by glass wool plugs. Molten wax was pumped from a heated graduated cylinder feed vessel along a heated line to the reactor inlet. Here it was mixed with hydrogen before entering and passing down the catalyst bed. Liquid and gaseous reactor effluents were first separated in a hot high pressure separator and then down stream in another high pressure separator cooled by carbon dioxide pellets. Off-gas from the cooled high pressure separator was vented through a rotameter. Careful

A simplified flow diagram for the Pilot 1 unit is given in Figure 1. The unit was a 20 year old pilot plant previously used for high pressure residue conversions that was modified for use in the FT wax hydrocracking studies. The design was similar t o the

"I

Gas

Pump

Figure 1. Pilot 1 u n i t experimental s y s t e m

Ri-HP hot separator, R2-LP sep., KO-cardice cooled sep.

hot separator, Kz-LP

refrig

4

John P. Collins et al./ Journal of Natural Gas Chemistry Vol. 15 No. 1 2006

MCE units but with an expanded product separation/product collection section and 3-zone tube furnace for temperature control. Here the catalyst charge was 100 ml of whole catalyst particles diluted 1:2 by volume with acid washed sand. The Pilot 1 unit suffered from a number of design and operability problems that adversely impacted the experimental results. As can be seen in Table 1, the reactor tube was relatively wide in diameter and short in length. The larger particle size (Dp) coupled with a wider reactor (d) and short length (1) meant a lower l / d and l / D , than the MCE units. This resulted in a higher Peclet number and increased axial dispersion. As a consequence, radial temperature profiles were typically 30 to 40 "C in the Pilot 1 unit compared to 3 "C in the MCE units. Frequent problems with the level control system required manual draining of the hot high pressure separator every 12 h instead of a continuous take off. Each manual draining caused a loss in pressure of typically 5 to 20 bar which disrupted steady reactor operation. Problems also occurred in quantifying the wax feed rate by the change of volume over time due to a density gradient in graduated feed cylinder. This coupled with poor control over the hydrogen feed rate resulted in low mass balances. The Pilot 2 unit (Figure 2 ) was designed to overcome the performance and operability problems previously encountered with the Pilot 1 unit, and to operate unattended at nights and on weekends. The most important deficiency of the Pilot 1 unit was its short bed length leading to poor hydrodynamics, high radial temperature differences, and lower middle distillate selectivities. A longer and narrower tube was used for the Pilot 2 reactor module which gave an l / D , value that was similar to the MCE units. The improved hydrodynamics reduced the radial temperature difference to only 5 "C. Temperature control was further improved by using a 6-zone tube furnace to control the reactor temperature. The accuracy of mass balances was improved by installing an electronic weight balance t o monitor the wax feed rate and hydrogen mass flow controller. For level control, a similar system was installed on the Pilot 2 unit but the differential pressure (dP) cell was duplicated and better heat tracing was provided for the cells and lines. These improvements enabled successful operation of the automated level control system. In order to meet the unattended operation requirements, a computer control system using proprietary software was incorporated for the Pilot 2 unit.

The software operated on a personal computer and carried out all control and data logging functions. Only the alarm and emergency shutdown system was hardwired. Control screens were utilized to adjust parameters such as temperatures, gas flow rates, levels, pumps, pressures and the on/off status of certain valves. Important operating parameters were trended on data logging screens. Such screens were extremely useful in trouble shooting during unattended operation since they showed exactly which control loop provoked a failure, and the subsequent sequence of events. Mass balance data were printed in tabular form daily, thus obviating the need for manual hourly readings.

Figure 2. Pilot plant 2 unit

4. Results of catalyst screening study Commercial hydrocracking catalysts covering a range of active metals, supports, and catalyst promoters were investigated in the screening studies (Table 2 ) . The base metal catalysts were operated in a nonsulfided state to further influence the functional balance toward a high middle distillate selectivity. The lack of sulfur in the feedstock and desire to exclude sulphur components from the product also made this a more favorable mode of operation. An assessment of Catalyst A operating in the sulfided mode was carried out for comparative purposes (Cat A Sulf.). The two Pt-alumina catalysts were available in chlorided form

5

Journal of Natural Gas Chemistry Vol. 15 No. 1 2006

(Catalyst C) or pretreated with CCl4 during catalyst loading (Catalyst D) to adjust the support acidity. The chloride is added to increase alumina acidity and improve the functional balance.

were probably too acidic, for the sulfur and nitrogen free paraffin feed. This was because both catalysts were designed for hydrocracking conventional crude oil materials, which normally contain nitrogen and sulfur species that reduce the hydrocracking activity.

Table 2. Catalysts tested in screening study ~~~~~~

Catalyst ID

Metal ( s )

Support

A B C D

Co, Mo Ni, Mo PtU

SiOz/A1203 SiOz/Alz03 A1203

chlorideb

P t U

A1203

chloride

E F

Ni Ni, Mo

SiOz SiOz zeolite-Y

A Sulf.d

Co, Mo

SiOz/AlzO?

100

Promotor

,

Cat. A Cat. B Cat. C o Cat. D A Cat. E A Cat. F v Cat. A-sulf. 8

n

T

8

8

0 8'

8

v 8 ' 8

en

q8'

8

sulfur"

a Catalyst C has a higher metal loading and pore volume than Catalyst D but lower support surface area, Catalyst also supplied pre-sulfided with 0.05 wt% sulfur, " Sulfided in situ, A Su1f.-Catalyst A in the sulfided mode.

"

0

"

~

20

0

"

~

~

60

40

"

80

"

100

Conversion (%)

Catalyst performance was accessed by comparison against the design case targets and through relative comparisons of activity, middle distillate selectivity, stability and product quality. The catalyst activities were expressed as the operating temperature required to achieve a given steady state conversion. Conversion is defined as the weight percent of >360 "C boiling materials in the wax feed converted to <360 "C while selectivity is the weight percent converted to the specified boiling range divided by the conversion. Selectivities to the gas oil and kerosene middle distillate fractions for each of the tested catalysts are plotted against conversion in Figures 3 and 4. The general trend for all catalysts was a decrease in the gas oil selectivity and increase in the kerosene and naphtha selectivities with increasing conversion. The screening results confirm the importance of the catalyst functional balance for optimal cracking of the ultra clean paraffin waxes. The most active catalysts (E and F) had the lowest selectivity to the desired middle distillate products. Catalyst F deactivated after the temperature was lowered in an attempt to reduce conversion and no products were collected. Synthetic crude from Catalyst E was pale yellow in color and barely mobile at ambient temperature containing large amounts of unconverted feed (n-paraffins). In addition, the cold flow properties of the gas oil fraction were poorer than those of materials derived from the other catalysts indicating a lower degree of isomerization within the gas oil range. The reason for these effects was due to neither catalyst possessing the optimal functional balance, both

Figure 3. Gas oil selectivity vs conversion during MCE screening studies

40

-5

t 30

1

-

.-2 . .-c al

2

20-

2

Cat. A Cat. B Cat. C o Cat. D A Cat. E A Cat. F v Cat. A-sulf. 8

0

P

8

8

In1

0

.

I-

o-s=j

'

I

'

'

'

'

'

'

'

I

'

'

'

1

'

'

'

1\

a

Catalyst C and D were the least active catalysts. Catalyst D was highly unstable under the operating conditions employed and rapidly deactivated providing no products. The operating temperature required to achieve a desired conversion level for Catalyst C was 40 to 50 "C higher than Catalysts A and B. Although selectivity was reasonable, the pour and cloud points for the gas oil fraction were high for a given conversion level. This indicates a higher concentration of waxy normal paraffins in the product caused by significant cracking out of the more highly branched paraffins at the higher operating temperature required.

"

"

6

John P. Collins et al./ Journal of Natural Gas Chemistry Vol. 15 No. 1 2006

Catalysts A and B exhibited good overall performance. Cetane indices for the gas oil products were high a t all conversion levels and the gas oil products had low cloud and pour points. The activity of Catalyst B was slightly higher reaching 60% conversion a t a 15 "C cooler temperature, but Catalyst A was more stable and provided a selectivity to the desired middle distillate products (Figure 2) that was superior to that of Catalyst B. Catalyst A was tested in both an unsulfided and sulfided state. The sulfided mode of operation would be employed when hydrocracking conventional crude oils containing sulfur species, but is not necessarily needed for upgrading the ultra clean wax produced in GTL conversion. The sulfided operating mode was achieved by a pretreatment step in which gas oil dosed with carbon disulfide was circulated over the catalyst under hydrogen. Following the pre-sulfiding treatment, wax dosed with 0.23 wt% dibenzyl-disulfide was fed to the reactor. The performance was better in terms of stability and selectivity in the unsulfided mode of operation. Catalyst A, a mild hydrocracking catalyst developed by BP and produced by a commercial catalyst manufacturer, was judged to have the best overall performance and was selected for further evaluation in micro-reactor and pilot units. The catalyst performance met the design standards established in 1990 and it produced gas oil products that were aromatics free with high cetane indexes and good cold flow properties. As shown in Table 2, it is a base metal (Co/Mo) catalyst manufactured from a silica alumina support. By the time the screening studies were conducted in the early 1990 s, it had already been used in conventional hydrocracking operations a t a BP refinery, in sulfided form, for over 20 years. 5. Further evaluations of selected catalyst Further evaluations were conducted to determine the stability and lifetime of the selected catalyst, effect of feed composition on hydrocracking performance, and the performance with whole catalyst particles in pilot units as opposed to the powdered version used in the MCE screening studies. An important factor leading t o the catalyst selection was the stability observed during 1000 h of screening tests. The stability was further demonstrated in the subsequent lifetime studies including MCE tests of 281 days and 150 days on-stream. The catalyst also proved to be robust against process upsets: loss in pressure, loss in wax feed, decreases in catalyst temperature, exotherms,

planned and unplanned shutdowns. In one case, it even recovered from a 4 h loss in the hydrogen feed during which the wax flow was uninterrupted. Based on the stability observed in these tests, a catalyst lifetime in excess of 2 years was predicted. 5.1. Wax feed effects The screening studies were conducted with a model wax feed obtained by blending two commercially available waxes. The resulting wax was lighter (average carbon number=35.6) than wax derived from the B P FT process (average carbon number=39.3). The effect of the feed composition was investigated using 4 waxes ranging from 30.9 t o 40.2 in average carbon number (molecular weight). Two B P FT waxes with different cut points and a commercially available FT wax were included in the investigation along with the model wax blend. It was observed that the higher molecular weight waxes were easier to hydrocrack than the lighter model wax which fits idealized paraffin hydrocracking theory [a] where adsorption and reaction of heavy molecules occur before lighter ones resulting in a greater reactivity for heavier feeds. Gas oil selectivity was also impacted by the carbon number distribution of the feed. Selectivities observed for the lighter model wax blend were higher than for the heavier FT waxes. The reason for the selectivity differences is related to the carbon number distribution of the feed and the mechanism for ideal hydrocracking [4]. A larger number of individual hydrocracking steps must occur before a higher molecular weight feed is converted t o <360 "C. Since the position of the carbon-carbon bond scission in each hydrocracking step is random, a wider distribution of product molecules is obtained resulting in a lower selectivity t o gas oil. These results indicate that selectivity patterns observed with the model wax feed are useful for showing trends and for catalyst performance comparisons but that further testing with the actual FT feedstock is desirable for absolute data. 5.2. Whole particle tests in pilot reactors

Good performance was achieved in the pilot reactor tests conducted with Catalyst A using both model and FT wax feed-stocks. The catalyst was again robust to hydrogen and wax feed failures, as well as loss in pressure confirming the stability observed in the MCE tests. Catalyst activities were approximately 5 to 10 "C higher in the pilot tests which can be ascribed

Journal of Natural Gas Chemistry Vol. 15 No. 1 2006

to increased liquid retentions over the catalyst for a given LHSV and less channeling whereby feed molecules pass through unconverted. Middle distillate selectivities in the Pilot 2 unit were similar to those observed in the MCE units, while Pilot 1 selectivities were generally 10 wt% less than expected. The poorer performance of the Pilot 1 unit was a consequence of poor hydrodynamics and the resulting high radial temperature gradients previously described. Therefore, the Pilot 1 unit may have suffered from high internal temperatures leading to local hot spots and over-cracking to gas and naphtha. 5.3. Akzo Nobel catalyst

By 1993 a hydrocracking catalyst for upgrading FT wax was selected, performance evaluated and economic studies performed. The technology was considered ready for progress to the demonstration scale but it was delayed. The delay was due to an unfavorable commercial environment at the time for GTL [4]. Another snag occurred, when the catalyst manufacturer ceased production of the selected wax upgrading cat-

7

alyst and a new one had to be found. A catalyst manufactured by Akzo Nobel (currently Albemarle) was identified as a possible replacement and testing commenced to prove its viability for wax upgrading in late 1996. The original catalyst and proposed replacement were tested in a recommissioned MCE unit using the BP FT wax as a feed. The Akzo Nobel catalyst proved to be more than an adequate replacement for converting the F T wax into synthetic crude. 6. The Nikiski GTL test plant

The sanction for BP’s GTL demonstration plant in Nikiski, Alaska, was given in July 2000. The plant is designed to produce a pumpable synthetic crude at 300 bpd scale. The integrated operation includes the compact reforming technology jointly developed by BP and Davy Process Technology, fixed bed FT Technology using a BP developed catalyst, and wax upgrading using an Azko Nobel hydrocracking catalyst selected in 1996. The synthetic crude is trucked off-site for further processing. An overhead view of the plant is provided in Figure 5.

Figure 5. GTL test Facility-Nikiski, Alaska

In the process flow scheme, wax feed is combined with a recirculating hydrogen stream, and heated to the desired feed temperature using a pre-heater and heat exchange against the hydrocracker product stream. The liquid product is separated from gas in hot and cold separators. Make up hydrogen is provided by the front end (syngas section) of the plant

and the gas purity is maintained by means of a small purge stream. The Nikiski plant underwent extensive commissioning and pre-start-up tests during 2002 and early 2003. Loading of the first product tanker truck on 21St July, 2003 (Figure 6) was a truly historic day in BP’s GTL program. Subsequent analysis confirmed

8

John P. Collins et aJ./ Journal of Natural Gas Chemistry VoJ. 15 No. 1 2006

that the high purity product contained no nitrogen, sulfur or aromatics. In the first plant start up, the hydrocracker was heated to cracking temperature by circulating diesel prior to wax introduction. An early milestone was achieved in the second start up by circulating the ultra-pure synthetic crude instead of diesel fuel to achieve hydrocracking temperature. Starting the unit up with the synthetic crude product minimizes the chance for wax blockages in the challenging climatic setting.

version while kerosene and naphtha selectivities increase. 700

,

I

600

Q 0

I ...................................................

0

10

20

30

40 50 60 70 Cumulative weight (‘33)

80

90

100

Figure 7. Boiling point distributions of hydrocracked product for different conversion levels (1) Conv.=57.5%, (2) Conv.=77.5%, (3) Conv.=96.25%

Figure 6. BP GTL team by first tanker of syncrude loaded at the site

A sampling of the test results is provided below. The hydrocracker has been operated at different conditions to access the impact of operating conditions on performance. Data from one set of tests are presented here. Table 3 summarizes the conversion of >360 “C boiling materials for three different operating conditions. The corresponding boiling point distribution curves of the product are shown in Figure 7. Since there was no adjustment to the boiling point distribution of the wax feed, it contains some components that boil below 360 “C. The low boiling materials are assumed to pass through unconverted and are not counted in the conversion calculation. The highest conversion was achieved at the lower wax feed flow even though the average temperature was a few degrees lower. The selectivity pattern was similar to what was observed in the early laboratory tests. The gas oil/diesel selectivity decreases with increasing conTable 3. Nikiski unit conversions at different wax feed flows (X in lb/hr, value between 600 and 2000) and temperatures (Y in degree F, value between 600 and 800) Relative

Relative

Conversion

temDerature

liauid flow

(wt%)

Y

X

96.25

Y+5 O F Y+10 O F

1.66X 1.66X

57.5 77.5

Progress has been steady since production of the first liquid product. The hydrocracking section is easy to operate and the catalyst has demonstrated a higher activity than in the earlier laboratory tests. Reasonable stability through extended operation, start ups and shutdowns has been observed confirming the results of the previous catalyst testing study. A large supply of wax produced in the FT convertor is now available for collaborative catalyst studies at Albemarle Catalysts Company BV. Use of the real FT wax will facilitate the ongoing catalyst studies and provide a better assessment of potential catalyst alternatives or variations in the operating conditions.

7. Engine emission tests Fractionation of the hydrocracked product from the two pilot plants produced sufficient material for diesel engine emission tests of F-T wax derived GTL diesel. The engine tests were conducted using GTL diesel as a neat fuel and in blends with conventional diesels. Initial tests were carried out in single cylinder research engines such as the Ricardo Hydra. The GTL diesel gave superior combustion and emissions performance under all conditions in these tests. The results suggested that when blended, the GTL diesel could be used to improve the performance of typical quality diesel fuel (Eurograde) to make a “premium” diesel with improved emissions performance. Alternatively, the GTL diesel could be blended with poorer quality refinery streams such as light cycle oil (LCO) to en-

9

Journal of Natural Gas Chemistry Vol. 15 No. 1 2006

able the more extensive use of poorer quality fuels in the diesel pool. A subsequent investigation was conducted with the multi cylinder Audi passenger car on the chassis dynamometer (rolling road emission tests) to determine whether benefits observed in the single cylinder engine tests could be obtained from a multicylinder engine under a legislative test procedure. Test fuels for the rolling road emission tests were as follows: 0 Fuel A Eurograde diesel fuel ex Rotterdam

(Base Fuel), 0 Fuel B GTL diesel, 0 Fuel C Eurograde diesel+GTL diesel (80:20 volume/volume) , 0 Fuel D Eurograde diesel+GTL diesel+LCO (60:25:15 volume/volume) The "Eurograde" diesel was representative of a typical quality European Diesel fuel (eg. 49 cetane number minimum). Properties of the various Fuels are listed in Table 4.

T a b l e 4. P r o p e r t i e s of test fuels used i n rolling road emissions s t u d y Parameter

Fuel A

Fuel B

Fuel C

Fuel D

Density at 15 "C (kg/L)

0.8404

0.7723

0.8270

0.8370

Kinematic viscosity at 40 "C (cSt) Kinematic viscosity at 50 "C (cSt)

0.322

2.36

2.99

3.01

2.65

1.98

2.47

2.49

Flash point ("C)

77

82

-

-

Sulphur content (%m)

0.195

0

0.168

0.220

Mono

13.7

0

11.5

13.3

Di-

5.0

0

4.7

7.4 2.3

Aromatics content (%vol)

Fuel Fuel Fuel Fuel

Tri-

0.7

0

1.0

Total

19.4

0

17.2

23

Cetane number

53.3

88

61.1

57.0

Cetane Index

54.9

60.1

56.5

A-Eurograde Base, B-GTL diesel, C-Eurograde diesel+GTL diesel (80:20 volume/volume), D-Eurograde diesel+GTL diesel+LCO (60:25:15 volume/volume)

The vehicle was fitted with a 1.9 liter direct injection engine equipped with a turbo charger, exhaust gas recirculation and an oxidation catalyst. The catalyst was sulfur tolerant so no difference in particulate emissions was expected due solely to the varying sulfur contents of the tested fuels. Full ECE-15.04 plus EUDC emission tests were carried out, measuring TEOM (tapered element oscillating microbalance) modal particulate data, and pre-catalyst, postcatalyst and tailpipe emissions. Averaged emissions from the rolling road tests are

reported in Table 5. Fuel D containing GTL diesel, Eurograde Diesel, and LCO had similar emissions performance t o the neat Eurograde diesel, Fuel A. The gaseous emissions were virtually the same while the blend gave slightly higher particulate emissions possibly due to its higher aromatics content. The premium blend of GTL diesel and Eurograde diesel (Fuel C) possessed a higher cetane number, and lower sulphur and aromatics content than the base fuel resulting in significant emissions benefits; gaseous emissions other than NO, were reduced as were particulate emissions.

T a b l e 5 . Averaged results f r o m a u d i rolling road emission tests Fuel

HC

Fuel A

0.438

Fuel B

0.133

Fuel C

0.317

Fuel D

0.427

co

NO,

co2

1.004

0.466

153.738

0.904

0.162

0.335

0.548

146.541

0.682

0.085

0.771

0.476

152.622

0.792

0.136

0.916

0.472

149.172

0.899

0.180

HC+NO,

Pm

All emissions reported in g/km Legislation 7/92 (new models) 31/12/92 (all new vehicles) C0<2.72 g/km, HC+N0,<0.97 g/km, P,<0.14 g/km HC-Hydrocarbon emissions, CO-Carbon monoxide emissions, NO,-Nitrogen oxide emissions, CO2-Carbon dioxide emissions, P,-Particulate matter

10

John P. Collins e t al./ Journal o f Natural Gas Chemistry Vol. 15 No. 1 2006

The neat GTL diesel produced the lowest emissions of all the fuels, except for NO, which was increased. This was not unexpected due t o the relatively high cetane number of the fuel. Particulate emissions were reduced by approximately 50% with the GTL diesel compared to the base Eurograde fuel. There were no reports of problems with driving the vehicle on blended or neat GTL diesel. The results indicate that GTL diesel can be blended with the base Eurograde fuel to produce a premium diesel with lower emissions. GTL diesel can also be used as a component to raise the quality of blends containing LCO to satisfy automotive diesel specifications without adversely affecting emissions performance.

References [l]Font Freide J H M, Gamlin T D, Graham C, Hensman J R, Nay B, Sharp C. Topics in Catalysis, 2003,

26(1-4): 3 [2] Conradt H L, Garwood W E. Ind Eng C h e m Process

Des Dev, 1964, 3(1):38 [3] Weisser 0, Landa S. Sulfided Catalysts, their properties and applications (ISBN 008017556 2). Oxford, New York: Pergamon press, 1973. 300 [4] Weitkamp J. Ind Eng C h e m Prod Res Dev, 1982, 21: 550 [5] Schulz H F, Weitkamp J H. Ind Eng C h e m Prod Res Dev, 1972, l l ( 1 ) : 46