Historical Developments in Hydroprocessing Bio-oils - Energy

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Energy & Fuels 2007, 21, 1792-1815

Historical Developments in Hydroprocessing Bio-oils Douglas C. Elliott* Pacific Northwest National Laboratory, P.O. Box 999, 902 Battelle BouleVard, Richland, Washington 99352 ReceiVed January 25, 2007. ReVised Manuscript ReceiVed March 16, 2007

This paper is a review of the developments in the field of catalytic hydroprocessing of biomass-derived liquefaction conversion products (bio-oil) over the past 25 years. Work has been underway, primarily in the U.S. and Europe, in catalytic hydrotreating and hydrocracking of bio-oil in both batch-fed and continuousflow bench-scale reactor systems. A range of heterogeneous catalyst materials have been tested, including conventional sulfided catalysts developed for petroleum hydroprocessing and precious metal catalysts. The important processing differences have been identified, which required adjustments to conventional hydroprocessing as applied to petroleum feedstocks. This application of hydroprocessing is seen as an extension of petroleum processing and system requirements are not far outside the range of conventional hydroprocessing. The technology is still under development but can play a significant role in supplementing increasingly expensive petroleum.

Introduction The objective of this review was to gather the relevant information from the literature (which is in specific cases more or less accessible from earlier times) to describe the development of hydrotreating upgrading technology to produce transportation liquid fuels from thermochemically derived biomass liquids. This paper provides only an overview of the types of technology for liquid fuels production from biomass, such as hydrothermal processing or fast pyrolysis. The important consideration is that these liquid products from biomass are not useful as fuels other than direct boiler firing and possibly for some types of turbine and large diesel applications after significant modifications. In order for the biomass liquids to be useful as transportation fuels, they require chemical transformation to increase volatility and thermal stability and reduce viscosity through oxygen removal and molecular weight reduction. The reader may refer to Bridgwater et al. for a discussion of how these types of technologies relate to each other.1 Maggi and Delmon provide a detailed comparison of fast and slow pyrolysis.2 The field of hydrotreating of hydrothermal (high-pressure) liquefaction oils from biomass provides the initial basis for this review. This study then focuses on the hydroprocessing of bio-oil from fast pyrolysis to liquid fuels. As such, it does not provide details of the liquefaction processes themselves but will focus on the catalytic processes used to convert the highly oxygenated biooil to a liquid more similar to petroleum-derived fuels. This review included collection of data from the literature in the field of study highlighting developments from the 1980s to the present. A useful summary of the early work has been published.3 This review is a collection and summary of the actual * E-mail address: [email protected]. (1) Bridgwater, A. V.; Meier, D.; Radlein, D. An Overview of Fast Pyrolysis of Biomass. Org. Geochem. 1999, 30, 1479-1493. (2) Maggi, R.; Delmon, B. Comparison between “Slow” and “Flash” Pyrolysis Oils from Biomass. Fuel 1994, 73 (5), 671-677. (3) Elliott, D. C.; Beckman, D.; Bridgwater, A. V.; Diebold, J. P.; Gevert, S. B.; Solantausta, Y. Developments in Direct Thermochemical Liquefaction of Biomass: 1983-1990. Energy Fuels 1991, 5 (3), 399-410.

data from the experimentation of handling and upgrading biooils by catalytic hydrotreating. Biomass Liquefaction Processes By way of introduction, the thermochemical conversion methods for producing oil from biomass are reviewed. There are basically two methodsshigh-pressure liquefaction (typically hydrothermal for biomass) and atmospheric-pressure fast pyrolysis. Hydrothermal Processes. As early as the 1920s, experimental data from Berl supported the concept of oil production from biomass in hot water using alkali as catalyst.4 Heinemann continued this work into the 1950s.5 Following the Arab oil embargo of 1974, initial efforts in biomass liquefaction in the U.S. focused on the high-pressure (hydrothermal) liquefaction process based on this concept developed at the Pittsburgh Energy Research Center (PERC) and demonstrated at the Albany Biomass Liquefaction Experimental Facility at Albany, Oregon,6 at a scale of 100 kg/h. In this process, wood powder was slurried in recycle oil product and mixed with water and a sodium carbonate catalyst. The slurry was pumped to a pressurized reactor where carbon monoxide was added at 20.8 MPa. The slurry was held at a temperature of around 350 °C for 20 min to 1 h after which it was depressurized, and oil product was separated from water. This heavy oil product had a melting point near room temperature with an oxygen content of 12 to 14% and dissolved water of 3 to 5%. A later derivative of this process was the HTU process from Shell.7 In this process, similar conditions were used but the (4) Berl, E. Science 1944, 99, 309. (5) Heinemann, H. Hydrocarbons from Cellulosic Wastes. Pet. Refiner 1954, 33 (7), 161-163. (6) Thigpen, P. L.; Berry, W. L. In Energy from Biomass and Wastes VI; Klass, D.L., Ed.; Institute of Gas Technology: Chicago, 1982; p 1057. (7) Goudriaan, F.; van de Beld, B.; Boerefijn, F. R.; Bos, G. M.; Naber, J. E.; van der Wal, S.; Zeevalkink, J. A. Thermal Efficiency of the HTU Process for Biomass Liquefaction. In Progress in Thermochemical Biomass ConVersion; Bridgwater, A.V., Ed.; Blackwell Science, LTD.: Oxford, England, 2001; pp 1312-1325.

10.1021/ef070044u CCC: $37.00 © 2007 American Chemical Society Published on Web 05/02/2007

DeVelopments in Hydroprocessing Bio-oils

reducing agent alkali/CO was left out; the resulting product had an oxygen content nearer to 18% with a melting point of about 80 °C.8 This process has been scaled up to a 10 kg/h dry biomass feed pilot plant. The pilot plant was operated for a 500 h design run in 2004. Fast Pyrolysis Processes. Most of the work up to the year 2000 is covered in detail in a review article by Bridgwater and Peacocke,9 but some of the highlights are described below. Modern developments in fast pyrolysis can be traced to the Occidental/Garrett process development work in the 1970s. That effort focused on waste processing, but a detailed report exists which describes results with biomass materials.10 The University of Waterloo first published results in a fluidized bed for fast pyrolysis of biomass in a government report in 1981. After extensive research, the technology was scaled up in Spain (Union Fenosa) and is now a patented technology held by Dynamotive Energy Systems Corporation. A company which grew out of the University of Waterloo, Resource Transforms International Ltd., is still developing refinements and chemical products from fast pyrolysis. A second version of fast pyrolysis which utilizes circulating fluidized beds was developed out of the University of Western Ontario and is now commercialized by Ensyn Technologies (RTP, rapid thermal processing). Their technology is used in food flavoring production plants in the U.S. It was scaled up for further development in Italy in the 1990s but was not operated extensively. Other versions of fast pyrolysis include the ablative reactor first demonstrated in its vortex form by the National Renewable Energy Laboratory (formerly the Solar Energy Research Institute) and now being developed at Aston University and Twente University with rotating plate and cone reactors. The Twente version has been scaled up in Malaysia. The simple entrained flow reactor was attempted at the Georgia Tech Research Institute in the 1980s as an extrapolation from the Tech-Air upflow pyrolysis/gasification technology. The operations could never achieve as high oil yields as those found in the other reactors, apparently because of the limited heat transfer. Vacuum pyrolysis was developed to demonstration scale at Laval University but was not found to be economical with biomass. Bio-oil Hydroprocessing Upgrading biomass-derived oils to hydrocarbon fuels requires oxygen removal and molecular weight reduction. As a result, there is typically a formation of an oil phase product and a separate aqueous phase product by hydroprocessing. To minimize hydrogen consumption in hydroprocessing, hydrodeoxygenation (HDO) must be emphasized, without saturation of the aromatic rings. Hydroprocessing of biomass-derived oils differs from processing petroleum or coal liquids because of the importance of deoxygenation, as opposed to nitrogen or sulfur removal. At the time that research began to evaluate the HDO of biomass-derived oils, HDO had received only limited attention in the literature.11 Over the past 20 plus years, there (8) Goudriaan, F.; Zeevalkink, J. A.; Naber, J. E. HTU Process Design and Development: Innovation Involves Many Disciplines. In Science in Thermal and Chemical Biomass ConVersion; Bridgwater, A. V., Boocock, D. G. B., Eds.; CPL Press: Newbury Berks, U.K., 2006; pp 1069-1081. (9) Bridgwater, A. V.; Peacocke, G. V. C. Fast Pyrolysis Processes for Biomass. Renewable Sustainable Energy ReV. 2000, 4, 1-73. (10) Boucher, F. B.; Knell, E. W.; Preston, G. T.; Mallan, G. M. Pyrolysis of Industrial Wastes for Oil and ActiVated Carbon RecoVery; report no. EPA-600/2-77-091, project no. S-801202, U.S. EPA: Washington, D.C., May 1977. (11) Furimsky, E. Catal. ReV.sSci. Eng. 1983, 25 (3), 42.

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have been wide-ranging efforts reported in the literature, as a recent review describes.12 A large portion of the body of work addresses the catalytic chemistry by hydrotreating model compounds containing oxygen. Many of these models are relevant to bio-oil hydroprocessing, such as phenolics and aromatic ethers. This report provides a detailed review of those research efforts that focused on processing actual bio-oil products, and model compound studies are addressed only when performed as a part of the bio-oil upgrading research effort. Hydrothermal Liquefaction Product Upgrading As the first wood liquefaction oil produced at pilot scale became available in 1979, oil upgrading process research began. The catalytic hydrotreating of the heavy oil product followed the methods used in conventional petroleum processing technology. Pacific Northwest National Laboratory (PNL/PNNL). Biooil upgrading work at PNNL has focused on heterogeneous catalytic hydroprocessing. Initial work involved batch reactor tests of model phenolic compounds13 with various catalysts. Commercial samples of catalysts were used representing CoMo, NiMo, NiW, Ni, Co, Pd, and CuCrO to hydrogenate phenol at 300 or 400 °C (1 h at temperature). P-Cresol, ethyl-phenol, dimethyl-phenol, trimethyl-phenol, naphthol, and guaiacol (methoxy-phenol) were also tested with a CoMo catalyst at 400 °C. Of the catalysts tested, the sulfided form of CoMo was most active, producing a product containing 33.8% benzene and 3.6% cyclohexane at 400 °C. The Ni catalyst was also active, producing a product with 16.9% benzene and 7.6% cyclohexane. The sulfided Ni catalyst still produced 8.0% cyclohexane, but its yield of benzene dropped to near zero (0.4%). A Pd catalyst produced a 7.8% benzene product with 2.7% cyclohexane but 5.5% cyclohexanone. At lower temperature (300 °C), cyclohexanone was the primary product at 8.1% and benzene and cyclohexane were nearly equal at 2.0 and 2.5%, respectively. The original work with hydrotreating biomass-derived liquids was the effort to make gasoline from the high-pressure liquefaction oil produced from wood at the Albany Biomass Liquefaction Pilot Plant. The oil from high-pressure liquefaction (a hydrothermal, alkali-catalyzed process) was a more deoxygenated product than fast pyrolysis bio-oil. As a result, it was more thermally stable and required less hydrogenation to produce a gasoline product. Processing technology was adapted from conventional petroleum hydrotreating using nickel and sulfided cobalt-molybdenum catalysts14 in a continuous-flow, fixed catalyst bed reactor system operated in an upflow configuration. The catalysts used in these tests are identified and described in Table 1. Preliminary results with a light oil fraction showed that the sulfided form of the CoMo catalyst was much more active than the oxide form. A copper-chromite catalyst was also tested and found to be even less active. The results are presented in Table 2 for extinction hydrotreating/hydrocracking of whole wood oil and a distillate product. The whole oil could be hydrotreated much like the distillate oil but required a lower space velocity and operating pressure and resulted in higher hydrogen consumption. The nickel catalyst exhibited activity (12) Furimsky, E. Catalytic Hydrodeoxygenation. Appl. Catal., A: Gen. 2000, 199, 147-190. (13) Elliott, D. C. Hydrodeoxygenation of Phenolic Components of Wood-Derived Oil. Prepr. Pap.sAm. Chem. Soc., DiV. Pet. Chem. 1983, 28 (3), 667-674. (14) Elliott, D. C.; Baker, E. G. Upgrading Biomass Liquefaction Products through Hydrodeoxygenation. Biotechnol. Bioeng. Symp. 1984, 14, 159-174.

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Table 1. Catalysts Used in Hydrotreating/Hydrocracking Tests supplier

catalyst id

active metals

weight %

support

forma

Harshaw Harshaw Harshaw Harshaw Harshaw Harshaw Haldor Topsoeb Haldor Topsoeb Haldor Topsoeb Haldor Topsoe Katalco Katalco Katalco Shell PNL/Linde PNL/Grace Amoco BASF Akzo Akzo Criterion Strem

Ni-1404 CoMo0402 HT 400 HT 500 Ni-3266 Ni-4301 TK 710 TK 750 TK 770 TK 751 CoMo479 CoMo499 KAT 4000 S411 CoMo/Y CoMo/SiAl NiMo/Y K8-11 KF-742 KF-840 C-424 78-166

Ni CoO/MoO3 CoO/MoO3 NiO/MoO3 Ni NiO/WO2 CoO/MoO3 CoO/MoO3 CoO/MoO3 NiO/MoO3 CoO/MoO3 CoO/MoO3 CoO/MoO3 NiO/MoO3 CoO/MoO3 CoO/MoO3 NiO/MoO3 CoO/MoO3 CoO/MoO3 NiO/MoO3 NiO/MoO3 Pt

68 3/15 3/15 3.5/15.5 50 6/19 2/6 2.3/10 3.4/14 3/13 4.4/19 4.4/19 3.5/14 2.67/14.48 3.5/13.9 3/13 3.5/18 4.3/11 4.4/15.0 3.9/19.6 4/19.5 5

proprietary silica-alumina γ-Al2O3 γ-Al2O3 silica-alumina γ-Al2O3 Al2O3 Al2O3 Al2O3 Al2O3 Al2O3 γ-Al2O3 γ-Al2O3 Al2O3 c Y-zeolite/ Al2O3 13%Al2O3 SiO2 Y-zeolite/ Al2O3 c MgO spinel/ Al2O3 γ-Al2O3 Al2O3c Al2O3c γ-Al2O3

1/8-in T 1/8-in T 1/8-in E 1/8-in E 1/16-in E 3/16-in R 1/16-in E 1/16-in E 1 mm E 1/16-in E 1/16-in E 1/32-in E 1/20-in T 1/16-in E 3/16-in T 1/16-in E 1 mm E 1.3 mm Q 1.3 mm Q 1.3 mm Q P

a E ) extrudates; R ) ring; T ) tablet; Q ) quadrilobe extrudates; P ) powder; and the size given is the o.d. b All three catalysts were used in a layered bed. c Includes phosphorus oxide.

Table 2. Hydrotreating Hydrothermal Products Experimental Operating Conditions catalyst CoMo 0402/S HT 400/S feedstock distillate whole oil temperature, °C 344 345 pressure, MPa 13.6 16.8 oil feed rate, mL/h 342 204 H2 feed rate, l/h 137 120 liquid hourly space velocity 0.38 0.24 (LHSV), vol oil/(vol cat h) H2 consumption, L/L oil 178 300 Experimental Results and Product Analyses carbon conversion, wt % to liquid product 70 77 to gas (C1 to C4) 1.0 2 to carbon on catalyst 12a 8 H/C atomic ratio oxygen, wt % moisture content, wt % density, g/mL sulfur in product, ppm

Wet Product Analysis 1.72 2.2 NA 0.89 NA

1.75 0.8 NA 0.82 441

Table 3. Hydrotreating Hydrothermal Products

Ni-1404 whole oil 333 16.3 274 300 0.30 538

69 15 9 1.82 3.5-6.1 1.7-2.9 0.89 NA

catalyst feedstock temperature, °C pressure, MPa LHSV, vol oil/vol cat/h H2 consumption, L/L oil

HT 400/S TR7 398 13.8 0.10 616

HT 400/S TR12 397 14.0 0.11 548

Yields total oil product, L/L feed aqueous phase, L/L feed C5-225 °C, L/L feed gas (C1 to C4), wt % carbon in aqueous H/C atomic ratio oxygen, wt % density, g/mL C5-225 °C, vol %

0.99 0.20 >0.86 14.1 0.1

0.92 0.20 0.34 9.0 NA

Wet Product Analysis 1.65 0.0 0.84 >87

1.5 0.8 0.91 37

similar to the sulfided CoMo catalyst except for having a much higher gas yield and requiring hydrogen consumption. The Ni catalyst seemed to lose activity over the several hours of operation as shown by the two product analyses in Table 2. Distillations of these products showed that they were mostly gasoline range material (50-225 °C BP). GC-MS analysis showed that the cyclic ketones and phenolics in the liquefaction product were converted to cyclic alkanes and aromatics. In further studies,15 comparisons were made between the whole oil presented above (TR7 water slurry, single pass) and an alternative form (TR12 produced with oil recycle). By GCMS analysis of the two feed oils, it was found that the TR7 oil contained primarily cyclic ketones and single-ring phenolics,

while the TR12 contained more multiring phenolics. Under similar processing conditions as shown in Table 3, a lighter hydrocarbon product was produced from the TR7 oil. A need for additional hydrocracking function in the catalyst was indicated for use with TR12. The relationships of space velocity to deoxygenation and gasoline yield when using the HT400-S catalyst are shown in Figures 1 and 2 for the two high-pressure liquefaction bio-oils.16 Catalyst stability was also evaluated and neither sulfur loss nor carbon buildup was judged to be the cause. The alkali content of the TR12 oil and its deposition on the catalyst was concluded to be the cause of catalyst deactivation by a pore plugging mechanism over a 48 h test. The resulting loss of activity in the mixed Haldor Topsoe catalyst system is shown in Figure 3. Extensive studies with several catalysts provided a useful set of results for comparing residual oxygen content and hydrogen to carbon ratio as well as gasoline yield as a function of space velocity16 (see Figures 4 and 5). It was concluded that high yields of high-quality gasoline can be produced from biomass-derived oils; however, low space

(15) Baker, E. G.; Elliott, D. C. Catalytic Upgrading of Biomass Pyrolysis Oils. In Research in Thermochemical Biomass ConVersion; Bridgwater, A. V., Kuester, J. L., Eds.; Elsevier Science Publishers, LTD.: Barking, England, 1988; pp 883-895.

(16) Baker, E. G.; Elliott, D. C. Catalytic Hydrotreating of BiomassDerived Oils. In Pyrolysis Oils from Biomass: Producing, Analyzing and Upgrading; ACS Symposium Series 376; Soltes, E. J., Milne, T. A., Eds.; American Chemical Society: Washington, D.C., 1988; pp 228-240.

a The distillate oil test was of limited duration, and the carbon on catalyst percentage was proportionally higher at 19%; but, the actual carbon weight percentage was similar to the other tests at 12%, suggesting that the carbon laydown occurs early in the run and reaches a steady state.



Figure 1. Oxygen content of hydrotreated oils (400 °C, 13.8 MPa).

Figure 2. Yield of gasoline boiling range (C5-225 °C) oil (400 °C, 13.8 MPa).

velocities are required. Cracking and hydrogenation of the higher molecular weight components are the rate-limiting steps in upgrading biomass-derived oils. Future catalyst development should be directed at these reactions. An advanced process concept was also tested wherein the initial hydrotreating for oxygen reduction was followed by separation of the aqueous, gasoline, and gas phases and a second processing step of hydrocracking the heavy components.17 By performing the multistep process, it was possible to

maximize the yield of aromatic gasoline by removing a fraction from the system before the rings became saturated. Furthermore, hydrogen consumption was minimized by using it only to hydrodeoxygenate the oxygen-containing components and it was not wasted on the saturation of aromatic compounds. (17) Baker, E. G.; Elliott, D. C. Method of Upgrading Oils Containing Hydroxyaromatic Hydrocarbon Compounds to Highly Aromatic Gasoline. U.S. Patent Number 5,180,868, January 19, 1993.


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Figure 3. Catalyst deactivation with Haldor Topsoe catalyst and TR12 Oil (400 °C, 13.8 MPa).

Figure 4. Effect of catalysts on the quality of hydrotreated oil from TR7 (400 °C, 13.8 MPa).

In batchwise recycle tests, the results of the recycle of heavy components to a hydrotreating bed (recycle concept) or to a hydrocracking environment (multistep concept) were compared. Both concepts showed good gasoline yield, but the hydrocracking catalysts were more active toward molecular weight reduction and produced more gasoline. These batchwise recycle tests were even more impressive when compared to the single-pass

test, as shown in Table 4. The two-stage hydrotreatment (with intermediate separation) results in a 13% reduction in hydrogen consumption for equivalent gasoline yield. The space velocity is increased by a factor of 4. With the use of a hydrocracking step, there was a similar reduction in hydrogen consumption and a similar fourfold increase in the space velocity resulting in a gasoline yield increase of up to 30%.



Figure 5. Effect of catalysts on the quality of hydrotreated oil from TR12 (400 °C, 13.8 MPa). Table 4. Comparison of Two-Step and Single-Step Resultsa two-stage H2 consumed, L/L LHSV, L/(L h) gasoline yield, L/L temperature, °C pressure, MPa a

single-stage

CoMo 499 (CoMo-HT in 2nd)

KAT 4000 (CoMo-HC in 2nd)

CoMo HT only

500 0.19 0.62 400/400 13.8

503 0.19 0.80 400/435 13.8

575 0.05 0.61 400 13.8

TR12 bio-oil feed was used in the HT step, distilled to remove volatiles 100 000 cps whose pour points would be around room temperature. These results suggest that only the highly upgraded oils with oxygen contents of 5% or less have the potential for direct use as turbine fuels because of viscosity limitations. Figure 9 also includes the data for the three bio-oils. The raw bio-oils show decreasing viscosity with increasing oxygen content because of increasing water content in the raw biocrude. Figure 10 shows the effect of temperature on the viscosity of several of the heavy oil products. These results suggest that the highly oxygenated products could not be used as turbine fuels because of their high viscosity even with preheating. The 10% oxygen oil might be used as turbine fuel with preheating to 50 °C or higher. Organic contamination of the water byproduct from hydrotreating is a concern relative to the overall wastewater treatment requirements for the plant. Figure 11 clearly shows that the contamination of the byproduct water as represented by the carbon content increases with the oxygen content of the



Figure 6. Effect of space velocity on product oxygen content.

Figure 7. Effect of space velocity on the yield of gasoline range hydrocarbon.

Figure 8. Processing rate effect on product oxygen content.

product oil up to a range of 5-10 wt % when the product oil oxygen content is >10%. At that level, the oil-water separation is difficult and oil phase contamination of the water samples leads to great variation in the results. Samples representing part and all of the oil product for two-stage upflow tests are given. The effect of biocrude feed properties on the reactivity can be significant. Whereas the work at PNNL over the years had used a number of bio-oils including several hardwood oils from different fluid-bed pyrolysis reactors, pine oil from ablative pyrolysis, hardwood oil from vacuum pyrolysis, and peat oil from fluid-bed pyrolysis without identifying any major differences in reactivity; more recent work showed large differences between poplar oil from ablative pyrolysis (NREL) and euca-

lyptus oil from fluid-bed pyrolysis (Union Fenosa). Table 14 provides some of the results. Although the eucalyptus-oil test was performed at higher temperature and lower space velocity, the product-oil quality was significantly lower, with higher oxygen content and density. Further substantiating the conclusion of lower reactivity is the lower gas yield and also the lower hydrogen consumption. The combination of high viscosity and poor pumping performance with the lower reactivity make the eucalyptus oil more difficult to hydrotreat. A number of comparisons were made between the PNNL work and that published by Veba Oel.39 The experimental reactor results are evaluated in terms of deoxygenation which incorporates the effects of temperature, pressure, and catalyst. The deoxygenation is defined as the percent of chemically combined oxygen in the bio-oil removed when compared to the oil product. Deoxygenation as a function of space velocity shows a dramatic effect. In Figure 12, space velocity is presented in terms of volume (liquid hourly space velocity). By this measure, the new downflow results in this paper stand out clearly compared to all of the earlier-published results. Thus, our earlier agreement with Veba that the effect of downflow operation compared to upflow was small considering the overall reaction and little difference between our earlier results and those (39) Baldauf, W.; Balfanz, U. Upgrading of Pyrolysis oils from Biomass in Existing Refinery Structures; final report JOUB-0015, Veba Oel AG: Gelsenkirchen, 1992.


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Figure 9. Relation of oil viscosity to oxygen content.

Figure 10. Effect of temperature on oil viscosity. Table 13. Low-Severity Hydrotreating Results in a Downflow Operation TK751 catalyst 1st temp, °C 2nd temp, °C LHSV, L/(L h) WHSV, g/(g h) yield, g/g biocrude deoxygenation, % density, g/mL gasification, % C H2 consumption, L/L carbon balance, % hydrogen carbon dioxide methane ethane propane butanes pentanes higher HCs

150 380 0.28 0.52 0.42 98.6 0.82 27 881 92

148 375 0.30 0.54 0.39 97.9 0.84 20 746 79

70.3 8.1 11.2 4.7 2.1 1.4 1.0 1.2

77.4 7.0 9.9 4.2 0.4 0.3 0.3 0.5

148 354 0.29 0.53 0.38 96.3 0.84 22 727 79

150 349 0.29 0.53 0.44 95.1 0.86 25 808 91

K8-11 150 349 0.29 0.54 0.49 95.0 0.86 33 813 107

148 362 0.43 0.78 0.47 94.5 0.86 29 791 101

Gas Effluent Composition (Including Excess Hydrogen) 78.4 70.7 72.2 73.3 5.8 7.6 5.5 5.5 8.0 9.8 7.5 7.1 3.9 4.8 3.8 3.5 1.2 2.2 3.6 3.5 0.9 1.8 2.6 2.4 0.8 1.5 2.2 2.0 1.0 1.6 2.5 2.6

of Veba. However, Figure 13 provides a data comparison on a weight basis (weight hourly space velocity) wherein the effect of the diluted CoMo catalyst bed used by Veba has a dramatic effect. By using the diluted CoMo catalyst bed, Veba achieved much higher processing rates based on the weight of catalyst. The one unexplained result is the poor showing of the NiMo catalyst in the downflow experiments at Veba. Again, our

148 355 0.38 0.70 0.53 95.8 NA 29 779 109 73.4 7.1 8.0 4.0 2.1 1.5 1.2 2.7

148 400 ( 15 0.38 0.64 0.21 97.6 NA 23 494 54 77.8 7.5 9.4 1.7 1.0 0.6 0.5 1.6

157 435 ( 45 0.31 0.51 0.21 97.1 NA 22 313 53 85.8 5.8 4.4 1.2 0.8 0.5 0.3 1.2

downflow results are much improved compared to our earlier upflow results. In all of these tests, the catalysts are supported on an alumina base. The differences in the use of CoMo versus NiMo are relatively small compared to the difference between upflow and downflow. The product-oil density is clearly a function of product-oil oxygen content as seen in Figure 14. The molecular weight may



Figure 11. Relation of carbon in the aqueous phase to the oxygen content of the oil. Table 14. Low-Severity Hydrotreating Results with Different Bio-oils temperature, °C WHSV, g oil/(g catalyst h) yield, g/g biocrude deoxygenation, % product density, g/mL gasification, % carbon hydrogen consumption, L/L oil carbon balance, %

NREL

Union Fenosa

355 0.70 0.53 96 0.86 29 779 109

365 0.54 0.41 92 0.94 17 554 85

also be correlated with the oxygen content as they may both be changing coincidentally as a function of processing severity. Figure 14 shows that the function of oxygen content and density seems to apply to the full range of products reported in the literature as well as those from these tests. An important conclusion to draw from these data is that the product-oil density approaches 1 at a relatively low oxygen content, about 10%. Products with oxygen contents around 10-15% and densities around 1 tend to form emulsions with the water byproduct and cannot be easily separated from the water. This lack of separation defeats one important purpose of the hydrotreating which is to remove the water and, thereby, dramatically improve the energy content of the oil. Veba Oel AG. Veba Oel performed a series of handling tests and upgrading experiments with biomass fast pyrolysis oil to evaluate its potential as a petroleum refinery feedstock Their initial assessment of the use of bio-oil in a petroleum refinery was evidently prepared before they had a useful understanding of the material, as they suggest either putting the bio-oil into a desalting unit or distillation columns.40 Neither option is reasonable, as the water-wash step, which is desalting, would dissolve more than half of the bio-oil into the water and the bio-oil would not distill in the fractionating columns but would polymerize to solids. In addition, the assessment for utilization (40) Rupp, M. Utilisation of Pyrolysis Liquids in Refineries. Biomass Pyrolysis Liquids Upgrading and Utilization; Bridgwater, A. V., Grassi, G., Eds.; Elsevier Applied Science: London, 1991; pp 219-225.

of bio-oil in a refinery draws the inappropriate conclusion that product from 50 pyrolysis plants would be required to displace only 5% of a conventional petroleum refinery. The conclusion is based on the faulty assumption that the pyrolysis plant would only operate at only 40 t/d of biomass feedstock whereas biomass conversion plant concepts are typically at least an order of magnitude larger and a single 2000 t/d plant, which is the concept used by the Department of Energy in the U.S., would provide the 5% displacement. After working with bio-oil,39 Veba’s conclusion was that, due to its immiscible nature, it was impossible to process bio-oil together with petroleum products without pretreatment. Following pretreatment, there are a number of possible sites where the upgraded bio-oil could enter the refinery process streams. They performed catalytic hydrotreating as the upgrading process to make bio-oil (RTP from Ensyn, filtered to remove the 2% char) compatible with petroleum. They used both sulfided CoMo and sulfided NiMo catalysts in a continuous-feed bench-scale reactor operated at 17.8 MPa. At weight hourly space velocities of 0.25-0.8 g/(g h) and temperatures of 350-370 °C, deoxygenation rates of 88-99.9% were achieved. Yields were relatively constant at 30-35% oil, 50-55% water, and 1520% gases. Hydrogen consumption ranged from 350 to 600 Nm3/t (420-720 L/L), 50-75% of which was used in water formation. At the high severity of 99.9% deoxygenation, the oil consisted of 40% naphtha, 40% middle distillates, and 20% vacuum-gas oil and was a yellow to white clear liquid. The heating value of the liquids was increased from 14.5 MJ/kg for the bio-oil to 40-44 MJ/kg for the upgraded oil. At severities below 95% deoxygenation, the product properties are unacceptable for introduction into the petroleum refinery.41 Veba reported that the catalyst deactivated quickly within 8 d. Gumlike deposits were found in the feeding tubes, valves, flow meters, and level controllers. At lower severity, the upgraded (41) Baldauf, W.; Balfanz, U. Upgrading of Fast Pyrolysis Liquids at Veba Oel AG. Biomass Gasification and PyrolysissState of the Art and Future Prospects. Kaltschmitt, M., Bridgwater, A. V., Eds.; CPL Press: Newbury, UK, 1997; pp 392-398.


Elliott

Figure 12. Deoxygenation on a liquid hourly space velocity basis.

Figure 13. Deoxygenation on a weight hourly space velocity basis.

Figure 14. Relation of product oil density to oxygen content.

oil became more and more soluble in water, which led to separation problems. With the CoMo catalyst, the reactor blocked in the preheating portion of the catalyst bed, and with NiMo, the coking was found in the outlet, not in the catalyst bed. A fixed-bed catalyst did not seem to be very practical and ebullated beds or liquid-phase systems (without catalyst) were suggested. Upgrading Tests. The upgrading tests were performed in a tubular reactor, 30 mm i.d. by 1.13 m long (0.716 L minus the cen-

tral thermowell), using both upflow and downflow configurations.42 The operating conditions tested and properties of some of the products are given in Table 15. Veba concluded that correlating with a decreasing deoxygenation rate there was an increase in specific gravity (see Figure 14), an increase in dissolved water (42) Baldauf, W.; Balfanz, U.; Rupp, M. Upgrading of Flash Pyrolysis Oil and Utilization in Refineries. Biomass Bioenergy 1994, 7 (1-6), 237244.



Figure 15. Relation of dissolved water content in the product oil to the level of deoxygenation. Table 15. Test Conditions and Properties of Upgraded Products catalyst flow mode temperature, °C space velocity, kg/(kg h) operating pressure, MPa chemical H2 consumption deoxygenation, % carbon hydrogen oxygen sulfur nitrogen density, g/mL water content, ppm CCR, wt % aromatic carbon, wt % HHV, MJ/kg LHV, MJ/kg 0-220 °C 200-350 °C 350-500 °C >500 °C

down 370 0.25 17.8 5.01 99.92

down 370 0.4 17.8 4.42 98.53

down 350 0.25 17.8 4.10 95.17

KF-702 down 350 0.4 17.8 3.68 91.78

down 370 0.8 17.8 3.16 88.32

Elemental Analysis, Water Free, wt % (Normalized to 100%) 86.79 87.22 86.84 86.5 85.94 13.17 12.35 11.82 11.23 10.79 0.021 0.39 1.25 2.14 3.06 0.07 0.004 0.004 0.008 0.055 0.003 0.033 0.082 0.121 0.152 0.83 0.86 0.896 0.928 0.94 50 210 510 1900 4200 0 0.09 1.40 3.39 5.52 10.0 17.0 18.8 25.9 27 45.96 45.26 44.46 43.43 41.2 43.09 42.6 41.87 40.99 38.87 40.8 38.5 20.7

40 36.2 21.2 2.6

Boiling Range, wt % 30.6 27.6 31.4 29.6 23.1 23.0 14.9 19.8

(see Figure 15), a decrease in heating value, an increase in Conradson carbon residue (CCR; see Figure 16), an increase in aromatic content (see Figure 17), and a decrease in gasoline and middle distillates while vacuum-gas oil remained fairly constant. The influence of the different catalysts and flow mode seemed to be as follows: • Under equal conditions, the deoxygenation in the upflow is lower than that in the downflow. • Deoxygenation with the NiMo catalyst is less than with the CoMo catalyst. • Depending on deoxygenation, lower amounts of naphtha are formed in upflow while the highest amount of vacuum gas oil is formed. • The aromatic carbon content seemed to be higher in upflow. • With the NiMo catalyst, there is a lower aromatic carbon content. Evidence of the catalyst deactivation is seen by comparison of data sets 2 and 6 at the same conditions but 150 h later, i.e., the deoxygenation has dropped from 98.53 to 95.88.

25.3 29.2 21.6 23.9

down 370 0.4 17.8 4.12 95.88

up 370 0.4 17.8 4.01 95.25

C424 down 370 0.1 17.8 5.01 97.20

86.93 11.87 1.06 0.005 0.129 0.90 350 1.79 18.8 44.38 41.68

86.82 11.52 1.2 0.007 0.23 0.931 351 1.60 22.7 43.44 40.91

86.47 12.90 0.6 0.021 0.009 0.866 102 0.28 9.8 44.83 42.00

29.8 30.2 23.8 16.2

21.4 32.1 27.6 18.9

38.5 36.3 22.3 3.0

Detailed analysis of the products was also undertaken. As shown in Table 16, several of the product fractions as well as the whole product oil were analyzed. From these analyses, it is concluded that the naphtha fraction has low octane because it contains high levels of cyclic compounds which will require reforming before it can be used in the gasoline pool at any more than small proportions; the diesel fraction is high density and low cetane because of a high aromatic content which will require hydrogenation to stabilize it; the vacuum-gas oil (VGO) meets the limits for CCR and metals content for use in hydrocracking; and the residue fraction is a very small portion which might not be separated in conventional distillation. It was recommended that the total product be routed to the crude distillation tower, where the fractions will end up diluted and upgraded in subsequent conversion units. Corrosion Tests. Corrosion problems were attributed to the high levels of organic acids in the bio-oil. Veba concluded that only special types of steels and packings would be suitable for processing plants. Corrosion tests were performed at 94 °C with


Elliott

Figure 16. Relation of Conradson carbon residue in the product oil to the level of deoxygenation.

Figure 17. Relation of aromatic content in the product oil to the level of deoxygenation.

the type 1.4541 FeCrNi-steel, the lowest grade steel used in the plants. The weight losses of the steel samples observed after 7 d of treatment amounted to 500 ppm with a corresponding rate of 70 ppm/d. The concentration of Fe, Cr, and, Ni increased from 69, 2, and 4 ppm in the bio-oil to 380, 210, and 47 after 7 d. At room temperature, no corrosion could be measured. During the upgrading tests, corrosion was observed on the thermowell (type 1.4570 steel). After changing to type 1.4980, an acid-resistant steel, there was less corrosion found. Stabilization Tests. Batch reactor hydrogenations of the biooil were performed at low temperature to evaluate the effect of a Pd catalyst for olefin removal. Through the use of NMR, a direct correlation of temperature from 25 to 155 °C with olefin removal was found; however, the overriding conclusion was that this type of processing led to a heavy bitumenlike product. Some water separation was also accomplished, but handling was difficult. However, Veba concluded that cheap alternatives for

stabilization and direct combustion and/or extraction of chemicals should be the focus of future activities. At that time (1994), the authors did not consider pyrolysis oil to be an economically attractive feedstock for a petroleum refinery, because the upgraded bio-oil cost would be from 300 to 800 ECU/t while petroleum and petroleum product costs were between 100 and 200 ECU/t. DMT FuelTec. A pilot-scale test of a slurry-phase hydrotreating technology was performed in 1997 by DMT.43 The Kohleoel pilot plant using the integrated gross oil refining (IGOR) technology was used for processing the heavy pyrolysis oil from the Union Fenosa plant. The bio-oil was processed with powdered NiMo catalyst at 30 MPa and 380 °C in a 2 m (43) Kaiser, M. Upgrading of Fast Pyrolysis Liquids by DMT. Biomass Gasification and PyrolysissState of the Art and Future Prospects; Kaltschmitt, M., Bridgwater, A. V., Eds.; CPL Press: Newbury, UK, 1997; pp 399-406.



Table 16. Characteristics of Hydrotreated Bio-oil Products

total product carbon, wt % hydrogen, wt % oxygen, wt % sulfur, ppm nitrogen, ppm calcium, ppm nickel, ppm vanadium, ppm molybdenum, ppm iron, ppm sodium, ppm density, g/mL n-paraffins wt % iso-paraffins, wt % naphthenes, wt % polynaphthenes, wt % aromatics, wt % monoaromatics, wt % diaromatics, wt % polycyclic aromatics, wt % MON/RON cetane number cetane index cloudpoint, C CCR, wt % stability test, ASTM D-2274

86.45 13.22 0.2 851 83

0.8503

naphtha

diesel

VGO

residue

500 °C

87.54 12.26 0.4 426 162

Historical Developments in Hydroprocessing Bio-oils - Energy

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