upgrading of flash pyrolysis oil and utilization in refineries

5/24/2018 Upgrading of Flash Pyrolysis Oil and Utilization in Refineries

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Pergamon

m w and B iocne rg~Vol. 7. Nos. l-6,pp. 237-244, 19940961-9534 94 000654 Copyright 1995 Elsevier Science iPrinted in Great Britain. All rights reserved961 9534194 $7.00 + 0.00

UPGRADING OF FLASH PYROLYSIS OIL AND

UTILIZATION IN REFINERIES

W. BALDAUF, U. BALFANZ and M. RUPP

VEBA OEL AG, Pawiker Str. 30, D-45876 Gelsenkirchen, Germany

Abstract-Flash pyrolysis oil from an ENSYN RTP pilot plant was upgraded in a continuous bench scaleunit with commercial CoMoand NiMocatalysts in anticipation of scaling up the process. Large amountsof product were produced in a pilot plant for use in an extended analytical characterisation programme.

In bench-scale experiments, high deoxygenation rates of 88-99.9 were achieved. Low liquid and highwater yields were obtained. The fractionated products of the production run did not fulfilthe requirementsfor direct use as gasoline and diesel.

The process is restricted by several operational problems such as rapid catalyst deactivation, coking and

plugging. Due to high feedstock and hydrogen addition costs, pyrolysis upgraded oil by the process testedis significantly more expensive than petroleum-derived oil at present oil prices.

INTRODUCTION

Pyrolysis oils from biomass are characterized bytheir high oxygen content. They are also veryunstable, and properties like viscosity and watercontent change during storage and heating. Theevidence for this is water formation and polym-erisation initiated by side reactions.

The properties and composition of pyrolysisoils mainly depend on the pyrolysis conditions.Rapid or fast pyrolysis processe? are charac-terised by a temperature range of 450-900Cand short residence times whilethe target of the reported project is to investi-gate possible routes for pyrolysis oils frombiomass to be processed in standard petroleumrefineries. Applied technology and experience

from the hydrotreating of conventional cokerdistillates are the background to this study.

PETROLEUM REFINERY

A petroleum refinery is a very complex se-quence of processes, which convert petroleumcrude oils into lighter marketable fuels andchemical feedstocks. A simplified scheme isshown in Fig. 1, which shows only the majorprocesses and major product streams.

The crude oil enters the refinery via thedesalting unit. After desalting the crude oil isatmospherically distilled and fractionated intonaphtha and middle distillate. The naphtha cut

is either routed to the naphtha pool, reformedto produce high-octane naphtha components, ortreated in olefin plants to produce ethene andpropene. The middle distillate can be directlyrouted to the middle distillate pool, or disul-

phurised prior to this. These pools supply gaso-line, jet fuel, diesel and fuel oil to the market.

The high boiling residue from the atmos-pheric distillation tower is further distilled in avacuum tower. The resulting distillates are cata-lytically cracked in hydrocrackers or fluid cat

crackers, respectively. The residues are ther-mally cracked in visbreaking or coking units, orin thermal hydrocracking processes. The lighter

products are treated in the same way as thestraight-run distillates.

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238 W. BALDAUF et al.

Atm.

isttll tton

v z.

iiill tton

Fig. 1. Refinery schematic.

Raw pyrolysis oils are not miscible with con-ventional petroleum crudes or any fractionsand this leads to the need for a pretreatmentstep prior to processing pyrolysis oil in therefinery.

FEEDSTOCK

Flash pyrolysis oil was provided by ENSYN(Canada). The oil was produced from mixedhardwood sawdust (maple/oak). The pyrolysis

plant operated at a temperature of 525C and aresidence time of 0.35 s. The ENSYN oil was

analysed with standard analytical methods.Results are listed in Table 1.

Depending on the quality of the pyrolysis oilafter pretreatment, it can be routed to eitheratmospheric or vacuum distillation, from whereit ends up very diluted in the three distillatestreams. Another possibility is to feed it directlyinto the conversion units. If pretreating productqualities allow, it could also be possible to distilthe hydrotreated oil separately and to feed thedistillates to the corresponding conversion unitsor to the corresponding product pools.

The possible feed points for pyrolysis oils areindicated in Fig. 1.

Table 1. Characterisationof raw pyrolysis oilWater content % 29.85Density (30C) g cm- 1.192HHV MJ kg- 16.23LHV MJ kg- 14.47Pourpoint C -30Flashpoint C 51Elemental analysis

(solids free)C wt 39.17H wt% 8.04S ppm 5

N wt% 0.050 wt% 52.74

Solids

Acetone insolubles wt% 2.2Pyridine insolubles wt 2.1Ash content wt% 0.15

Because the crude pyrolysis oil contained 2%of char the oil had to be filtered prior tohydrotreating.

EXPERIMENTAL

The upgrading test runs were performed ina continuously operating bench-scale catalysttesting unit. The unit is designed for pressures of

up to 30 mPa, temperatures of up to 500C andflow rates up to 300 g h-. Especially for pro-cessing pyrolysis oil the unit was modified for anoptional down- or upflow mode operation. Asimplified flow scheme of this continuous-flowfixed-bed unit is shown in Fig. 2.

The make-up hydrogen was supplied from ahigh-pressure hydrogen supply and the liquidfeedstock was fed by means of a high-pressurepump. Both feed rates were continuouslymeasured by flowmeters. After contacting gas

and liquid the reactor was charged in the se-lected mode.

The reactor (30mm internal diameter and1130 mm long) had a volume of 7 16 cm3 minusthe volume of a central thermowell which con-


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Upgrading of flash pyrolysis oil 239

Reactor

Stab. Gas

Fig. 2. Bench-scale unit for catalyst testing.

tained thermocouples for measuring thetemperatures along the reaction section. Thereactor was externally heated by five

heating zones which were independently con-trolled and monitored by means of the thermo-couples.

Preheating was performed in the first part ofthe reactor, which was filled with inert SIC. The

centre part of the reactor was filled with cata-lyst. The feedstock was further heated up to thereaction temperature in the first third of the

reactor which was isothermal in the remainingcatalyst zone. The catalysts were standard refin-ery catalysts in their sulphided form. To preventdesulphiding and thus deactivation of the cata-lyst small amounts of DMDS (dimethyldisul-


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240 W . BALDAUF et a l .

WHSV: 4

WHSV: 8

34 0 350 360 370

Temperature lFig. 3. Influence of temperature on deoxygenation.

380

phide) were added to the feed to ensure aminimum H,S partial pressure.

The product leaving the reactor was cooleddown at pressure. The offgas was separatedfrom the liquid phase in the high-pressure sepa-

rator and depressurised by a control valve whichestablished the operating pressure of the unit.The liquid phase was depressurised and separ-ated into a liquid product and stabilisation gasin the stabiliser.

The production run was performed in a hy-drotreating pilot plant which was designed foroil flow rates of up to 10 kg h-.The configur-ation of the plant is similar to that of thebench-scale unit described above. A detaileddescription of the plant is given in ref. [8].

Material balances were performed bysampling the liquid and gaseous products. Theobservation of mass balance was 48 h afterestablishment of stable conditions.

Raw mass balances were corrected by a calcu-lation procedure which closed elemental bal-ances of carbon, hydrogen, oxygen, sulphur andnitrogen. Hydrogen consumption was calcu-lated as the difference in hydrogen content ofthe product streams minus hydrogen content inthe liquid feed. H,S,NH3 and additional H,Owere calculated by balancing S, N and 0 inliquid product and feed (water in feed oil andDMDS were taken into account). Finally

flow rates were corrected by closing the carbonbalance. C,-C,,CO, CO*,and C,,were calcu-lated from corrected flow rates and the compo-sitions of liquid products, off gases and strippergases.

RESULTS

During the experiments the effects of tem-pe r a tu r e (350-37OC), s pa c e ve loc i t y(0.154.80 kg (kg h))), up- and downflowmode, and different catalysts (CoMo, NiMo)were investigated.

In order to reduce the number of independentvariables (temperature, space velocity) theseverity parameter deoxygenation rate wasintroduced for the following interpretation ofthe test results. Deoxygenation rate representsthe reduction of concentration of organic oxy-gen in product oil related to organic oxygenconcentration in the feed. Depending on theoperating conditions, deoxygenation rate in-creases with increasing severity which is illus-trated in Fig. 3.

The results of analytical inspection of product

oils are summarised in Table 2. The correspond-ing balances and product distributions are listedin Table 3.

Oil yields decrease from 36 to 32% withincreasing deoxygenation. Water shows thehighest yields. In addition to the water in thefeed, 21-25.5 wt% of water was produced. Gasand water yields increase at higher severities. Inaddition to water formation, significantamounts of oxygen were removed as CO andc o , . With increasing deoxygenation the

amounts of CO and CO*decrease. This meansthat, at higher severity, oxygen removal shiftsmore and more to water production instead ofCO, production. Oxygen rejection via waterincreases hydrogen consumption. As deoxy-


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Upgrading of flash pyrolysis oil

WV

243

0 50 104 150

Pyrolysis Oil ECU/t

Fig. 4. Economics of upgrading.

Total l i quid product

The product oil is a light brown low viscosityliquid with a boiling range from 35-535C.Napht ha (boi l i ng range 180C)

High naphthene content and low octane num-ber reduce its suitability as a gasoline poolcomponent. Only small amounts can be blendedas gasoline. Further upgrading in the reformerplant and conversion to olefins in steam crack-ers are more promising routes.

Di esel (boi l i ng range 180 -350C)The diesel fraction is characterised by high

density, low cetane number and high aromaticscontent. This fraction does not meet the specifi-cations of a marketable diesel product. Theoxygen stability test was negative. The need forfurther upgrading of this fraction for stabilis-ationand aromatics reduction prior to use as adiesel pool component is evident.

Vacuum gasoil{ VGO (boi l i ng range 350 -500C)CCR and metals content of VGO are at the

limits recommended for feedstocks for hydro-cracking (HC) and fluid cat cracking (FCC)plants. Alternatively, the fraction might be usedwithout further upgrading as heavy fuel oil.

Residue ( 500 C)

The residue fraction represents only a verysmall amount with a final boiling point of535C. Commercial vacuum distillation unitswill not separate this fraction because VGO cut

points usually are >535C. That means that theresidue will be part of the VGO fraction and istreated as described above. Restrictions mightoccur if CCR and metal contents of vacuum gasoil are significantly changed by the shift of thecut point.

Results of yield structure and fraction qual-

ities indicate that the total liquid product cannotbe fed to a single further upgrading process.Separate fractionation or feeding to the pet-roleum crude oil distillation unit (which is moreeconomic in an existing refinery) is necessary to

split the oil into the main fractions.The reported possible suitability routes are to

be evaluated on the basis of the major analyticalproperties of the fractions. The final decision forsuitability has to be confirmed by experimentaltest programmes.

ECONOMICS OF UPGRADING

Low yields and high hydrogen costs are re-sponsible for the high cost of upgraded pyrolysisoil. In Fig. 4 the costs for highly upgraded oilare shown as a function of raw pyrolysis oil

price. Product costs are calculated as the sum ofthe feedstock cost, upgrading cost and hydrogencost, minus a credit for the hydrocarbon gaseson the following basis:

??raw pyrolysis price is varied between 50 and15OECU t-;??upgrading costs are 15 ECU t- feedstock at a

hydrotreating plant capacity of0.8 x lo6t year-;

??hydrogen costs are 0.09 ECU Nmd3for steamreforming of natural gas;

??hydrocarbon product gases are credited at aprice proportional to their heating value whichleads to approx. 30-50 ECU t- of gas.The result is that upgraded pyrolysis oil is

more expensive by a factor of 1.5-3 than equiv-alent petroleum oil products. At typical feed-stock costs of 100 ECU t-for raw pyrolysis oil,1 t of upgraded oil costs about 450 ECU t-l,compared with 18&220 ECU for petroleum-


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244 W. BALDAUF et al.

derived naphtha and middle distillate andSO-100 ECU t- for fuel oil. The major contri-bution to this cost is the feedstock price (3 t offeedstock are necessary to produce 1 t of up-graded pyrolysis oil); about 150 ECU t- ofproduct has to be expended for hydrogen.Credits for the hydrocarbon product gases andthe upgrading costs offset each other.

The situation is similar for partially upgradedoil, compared to prices of conventional fuel oils.

The value of extenuities such as CO, tax or acredit for zero sulphur content can improve theprice relations for biomass resources. Thesebenefits are reduced if the hydrogen and processenergy production by fossil resources are takeninto account.

CONCLUSIONS

Catalytic upgrading of flash pyrolysis oil in afixed-bed reactor system was realised for shortrun times. High deoxygenation rates were ob-tained. Water is the major product and the oilyield is only 30-35 . The products are lightliquids with low oxygen and water content andhigh heating values. At lower severity of lowerdeoxygenation rate the solubility of product oil

in water is increased and the separation ofaqueous and organic phases is increasinglydifficult.

The catalyst deactivated very rapidly andsteady-state conditions could only be obtainedfor a short period. Several operational problemscaused by the instability and coking potential ofthe pyrolysis oil made processing difficult.Tubes and valves were plugged or covered withgum-like deposits. Operating and control sys-tems such as flow meters and level controllers

were disturbed by adhesive deposits. WithCoMo catalysts the catalyst bed plugged andcoked in the preheater, whereas with a NiMocatalyst the reactor outlet was plugged by gum-like substances.

Although the products are already hydrogen-ated, some fractions did not fulfil the required

product specifications.The naphtha fraction is unsuitable for direct

use as gasoline. Only very small amounts couldbe routed to the gasoline pool. An additionalupgrading (e.g. in a reformer) or the use of thenaphtha fraction as a feedstock for olefin pro-duction should be more promising.

Low cetane number (45) high density(0.873 g ml-), and oxygen instability illustratethe poor potential of the diesel fraction fordirect use as a pool component. Further upgrad-ing is necessary.

Total liquid product from the pretreatmentstep is recommended to be routed to the crudedistillation tower, where the fractions end updiluted in the naphtha, diesel and VGO cut andare further upgraded in subsequent conversionunits together with petroleum-derived products.

A fixed-bed reactor system is not recom-mended, since the rapid catalyst deactivationmust be compensated by continuous regener-ation. An ebullated bed or liquid-phase reactorsystem with homogenous or without catalysts

might be more promising.Acknowledgements-The work was performed with thefinancial support of the European Union, DG XII in theframework of the JOULE programme. We thank ENSYN,Canada and Energy and Mines Resources, Canada for

providing the pyrolysis oil.

REFERENCES

M. Polk and M. Phingbodhipakkiya, U.S. Environmen-tal Protection Agency, Off. Res. Dev. [Rep.] EPA,EPA-600/9-79-023b, Munic, Solid Waste: Resour. Re-covery, pp. 118-125 (1979).

2. R. G. Graham, B.A. Free1 and M. A. Bergougnou, Theproduction of pyrolytic liquids, gas and char from woodand cellulose by fast pyrolysis, In Research in Thermo-chemical B iomass Conversion (eds A. Bridgwater andJ. Kiister), p. 629. Elsevier Applied Science, Londonand New York (1988).

3. R. G. Graham, B. A. Freel, D. R. Huffman and M. A.Bergougnou, The production of liquid fuels and chemi-cals from biomass by rapid thermal processing (RTP),1st European Forum on E lectri city Production fr omBiomass and Solid Wasresby Advanced Technologies,Florence, Italy (27-29 November 1991).

4. D. Scott and J. Piskorz, The flash pyrolysis of as-pen-poplar wood. Can. J Chem. Engng 60,666 (1982).

5. R. G. Graham, 8. A. Freel, D. R. Huffman and A. J.Vogiatzis, The characterization and combustion of fast

pyrolysis bio-oils, 1st European Forum on ElectricityProduction fr om Biomass and Solid Wastes by Advanced

Technologies, Florence, Italy (27-29 November 1991).6. E. G. Baker and D. L. Elliott, Catalytic upgrading of

biomass pyrolysis oils, In Research in ThermochemicalBiomass Conversion (eds A. Bridgwater and J. Kiister),p. 883. Elsevier Applied Science, London and NewYork (1988). __

7. E. Churin, R. Maggi, P. Grange and B. Delmon,Characterization and upgrading of a bio-oil produced

by pyrolysis of biomass, InResearch in ThermochemicalBiomass Conversion (eds A. Bridawaterand J. Kiister).p. 896. Elsevier Applied Science, London and NewYork (1988).

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upgrading of flash pyrolysis oil and utilization in refineries

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