US3363024A

US3363024A - Hydrogenation of hydrocarbon oils

Info

Publication number: US3363024A
Application number: US605503A
Authority: US
Inventors: Majumdar Binay Bhushan; Figueiredo Osmond; Murthy Panchagnula Srinivasa
Original assignee: Gas Council
Current assignee: Gas Council
Priority date: 1966-12-28
Filing date: 1966-12-28
Publication date: 1968-01-09
Anticipated expiration: 1985-01-09

Description

1963 B. a. MAJUMDAR ETAL 3,363,024

HYDROGENATION OF HYDROCARBON OILS v Original Filed March 8. 1963 .lnivENToRs 5. B. Majumdar' O. Fisueiredo F. S. Mar-Hwy x ,ZM

ATTORNEY United States Patent HYDROGENATION 0F HYDROCARBON OILS Binay Bhushan Majumdar, Solihull, Osmond Figueiredo,

Birmingham, England, and Panchagnula Srinivasa Murthy, Bombay, India, assignors to The Gas Council,

London, England, a British statutory corporation Continuation of application Ser. No. 263,930, Mar. 8,

1963. This application Dec. 28, 1966, Ser. No. 605,503

19 Claims. (Cl. 260-6833) ABSTRACT OF THE DISCLOSURE The aliphatic constituents of a hydrocarbon distillate oil are converted into methane or methane and ethane by the exothermic reaction of the vapour of the oil with hydrogen at 600 to 800 C. under 20 atmospheres pressure. A stream of the reactants is continuously introduced into a thermally insulated reaction chamber to cause a substantial body of gas comprising reactants and reaction products to circulate continuously in the chamber. The reactants are so preheated as to maintain in the chamber a reaction temperature that is substantially uniform throughout the chamber, and gaseous products of reaction are continuously withdrawn from the chamber.

This application is a continuation of Ser. No. 263,930 filed Mar. 8, 1963, and now abandoned.

This invention relates to the vapour-phase hydrogenation of hydrocarbon distillate oils comprising aliphatic constituents.

In the process of this invention the oil vapour is reacted exothermically with hydrogen at a temperature within the range of from 600 C. to 800 C. under a pressure of at least 5 atmospheres gauge. The process results in the conversion of substantially the whole of the aliphatic constituents of the oil into methane or a mixture of methane and ethane (methane and mixtures of methane and ethane being hereinafter referred to generally as saturated hydrocarbon gas), and with the proportion of hydrogen required to bring about the said conversion the overall reaction is exothermic.

The expression hydrocarbon distillate oils comprising aliphatic constituents is used in this specification to mean oils that are distillates and consist wholly or predominantly of hydrocarbons, and of which the aliphatic constituents are aliphatic hydrocarbons and/or the aliphatic portions, such as side chains, of molecules of alkylated aromatic hydrocarbons. The oils, being distillates, contain substantially no non-vaporisable matter. The term hydrocarbons is used herein to include both unsubstituted hydrocarbons and hydroxy-substitution products of hydrocarbons of which the hydroxyl groups are split off during the reaction.

The hydrocarbon distillate oils may be petroleum distillates, advantageously light distillates having final boiling points within the range of from 70 C. to 200 C., for example, a light distillate having a boiling range of 32 C. to 170 C.; but they may be heavier distillates, for example, kerosene having a boiling range of 160 C. to 285 C. or gas oil having a boiling range of 180 C. to 360 C., or they may be liquefied gases that consist mainly of butane and/ or propane and are obtained in the distillation of petroleum. Hydrocarbon distillates not obtained from petroleum may also be used, such as coal tar distillates containing unsubstituted hydrocarbons and/or hydroxy-substitution products of hydrocarbons, for example, creosote oils.

The invention provides a continuous process for the vapour-phase hydrogenation of a hydrocarbon distillate Patented Jan. 9, 1968 oil comprising aliphatic constituents, wherein the oil vapour and a gas comprising hydrogen are continuously introduced in the form of a jet or jets through orifice means into a thermally insulated reaction chamber in which the oil vapour reacts exothermically with hydrogen at a temperature within the range of from 600 C. to 800 C. under a pressure of at least 5 atmospheres gauge, gaseous products of reaction are continuously withdrawn from the reaction chamber, the reactants are introduced through the orifice means at a high velocity and the arrangement is such as to cause a substantial body of gas comprising both reactants and reaction products to circulate continuously Within the chamber, and the reactants are preheated to an extent such as to maintain a reaction tern perature within the aforesaid range throughout the interior of the reaction chamber except in the vicinity of the orifice means.

The hydrogenating gas, that is to say, the gas comprising hydrogen, advantageously consists substantially wholly of hydrogen, but it may be a gaseous mixture consisting mainly of hydrogen (measured by volume). Examples of such gaseous mixtures are mixtures that contain hydrogen and carbon monoxide and are obtained by the reaction of carbonaceous materials, for example, coal, coke or hydrocarbons with steam. A mixture so obtained may be converted into a gas consisting substantially wholly of hydrogen by reacting a part or the whole of the carbon monoxide therein with steam, in accordance with the known water gas switch reaction, to form hydrogen and carbon dioxide, and subsequently removing the bulk of the carbon dioxide.

For the reasons stated above, the fact that the overall reaction is exothermic implies that the proportion of hydrogen introduced by the hydrogenating gas relatively to the proportion of oil vapour is at least sulficient to convert substantially the whole of the aliphatic constituents of the oil into saturated hydrocarbon gas. The proportion of hydrogen so introduced will generally exceed the aforesaid minimum proportion. There is usually a tendency for carbon to be deposited on the wall of the reaction chamber so as ultimately to impair the recirculation. The proportion of hydrogen is an important factor governing the rate at which this deposition occurs, the rate increasing as the proportion of hydrogen is decreased towards the aforesaid minimum, assuming that the other relevant factors, Which are discussed hereinafter, remain unaltered. Accordingly, it generally is desirable to use an appreciable excess of hydrogen.

Since any unreacted hydrogen lowers the concentration of saturated hydrocarbon gas in the product gas, the introduction of an excess of hydrogen can also be utilised to reduce the calorific value of the product gas to a desired extent. In the case of light petroleum distillates the quantity of distillate vapour introduced with each 1000 cubic feet (measured at S.T.P.) of hydrogen in the hydrogenating gas is advantageously within the range of from 2 to 8 gallons (measured as the liquid distillate).

The rate of deposition of carbon can be markedly decreased by using a hydrogenating gas containing a suitable proportion of steam. Thus, in the treatment of light petroleum distillates, the presence of 10% by volume of steam in the hydrogenating gas has been found very advantageous for this purpose.

The reaction temperature is preferably within the range of from 700 C. to 800 C. in order to obtain a desirably high rate of reaction, but the rate of carbon deposition increases as the temperature increases.

In order to maintain a desired reaction temperature, the oil vapour and the hydrogenating gas will usually be both preheated and introduced through the orifice means in admixture with one another, but in some cases it may 3 be necessary or desirable to preheat them separately to different temperatures and to mix them together before or upon being so introduced. In all cases the oil vapour and the hydrogenating gas are so preheated that their mixed preheat temperature (that is to say, the temperature of the preheated mixture or the temperature the reactants have when mixed together) is below the reaction temperature, but is high enough to maintain the desired reaction temperature. Depending upon the heat of reaction, the rate at which the reaction proceeds, the effectiveness of the thermal insulation of the reaction chamber, and the internal dimensions of the chamber, the mixed preheat temperature may be, for example, within the range of from 400 C. to 550 C., or even below 400 C., for example, 350 C. Using a large reaction chamber having well insulated walls, so that the rate of loss of heat through the walls per unit internal volume of the reaction chamber is small, it is possible, when carrying out the hydrogenation of a light petroleum distillate having a final boiling point of 170 C. and using a hydrogenating gas containing at least about 90 to 95% by volume of hydrogen, to maintain a reaction temperature of 750 C. while introducing the reactants into the reaction chamber at a temperature of only 350 C.

The pressure within the reaction chamber is preferably within the range of from 20 to 50 atmospheres gauge. The rate of deposition of carbon decreases as the pressure is increased. The gaseous products of reaction are advantageously withdrawn from the reaction chamber through valve means arranged so to control the rate of withdrawal of the gaseous products of reaction as to maintain the pressure within the reaction chamber substantially constant.

The circulation of gas within the reaction chamber is caused by the transfer of momentum from the rapidly moving stream or streams of gas entering the reaction chamber to the gas already in the reaction chamber. The magnitude of the circulatory effect may be specified in terms of the recirculation ratio, that is to say, the ratio of the volume of gas circulating within the reaction chamber to the volume of gas withdrawn from the reaction chamber during one complete period of the circulatory motion.

The recirculation ratio may be at least 3:1, is advantageously at least :1 and preferably at least 20:1. In general, a high recirculation ratio is desirable because, especially when the degree of preheat is low, it decreases the period that elapses after the reactants are introduced into the reaction chamber and until they are brought up to a temperature, say, about 650 C. at which the aforesaid reaction proceeds rapidly, and this results in a marked decrease in the rate of carbon deposition. On the other hand, the conditions necessary to obtain a high recirculation ratio may, as is explained hereinafter, be disadvantageous from other points of view.

For applications in which some deposition of carbon is not a serious disadvantage, some guide as to whether or not the recirculation ratio is sufficiently high may be obtained by making temperature measurements, for example, by means of thermocouples, at various points within the interior of the reaction chamber. The temperatures so measured should, except in the immediate vicinity of the orifice means, all be uniform to within :10 C. and preferably to within :5" C. or even :25 C. This criterion is not, however, always sufiicient to ensure that the rate of deposition of carbon is sufliciently low and, when that consideration is of great importance (as, for example, when it is desired to operate the process continuously for a period of about one year), the rate of deposition of carbon should itself be measured.

The main factor that determines the magnitude of the recirculation ratio is the ratio of the cross-sectional area of the region of the reaction chamber into which the reactants are introduced to the cross-sectional area of the orifice means (each taken in a plane that is perpendicular to the direction of flow of the gases). Assuming that, as is desirable, the reaction chamber is so designed that the resistance offered by the reaction chamber to the circulatory flow of the gas is as low as possible, it will usually be found that the recirculation ratio is very roughly equal to half the square root of the aforesaid ratio of crosssectional areas, which may be at least :1 and is preferably at least 400:1. If a more accurate estimate of the recirculation ratio is required, this may be obtained by using a model of the reaction chamber, introducing a gas (conveniently air) through the orifice means and measuring the rate of recirculation of air within the chamber and the rate at which air flows out through the outlet (which, in the steady state, must be equal to the rate of supply of air to the orifice means). Another factor that affects the magnitude of the recirculation ratio is the ratio of the temperature of introduction of the reactants to the temperature of the gas within the reaction chamber (outside the immediate vicinity of the orifice means), the recirculation ratio being proportional to the square root of this ratio if the temperatures are expressed on the absolute scale. This latter factor is not usually significant unless the degree of preheat of the reactants is very low in relation to the temperature within the reaction chamber, but it is of importance in those circumstances, because the effect then lowers the recirculation ratio appreciably and, as is explained hereinbefore, in such circumstances the use of an unduly low recirculation ratio is to be avoided. It is also of importance, and must be allowed for, in the determination of the recirculation ratio by measurements made on a model using air and with no substantial difference in temperature between the air entering the chamber and the air already in the chamber.

For the reasons given hereinbefore, if it is desired to increase the recirculation ratio for a given reaction chamber, it is necessary (if any very considerable increase is to be achieved) to decrease the cross-sectional area of the orifice means and, if the throughput is not to be correspondingly reduced, this requires an increase in the velocity of introduction of the reactants into the reaction chamber, together with a corresponding increase in the pressure drop across the orifice means and in the power consumption. Assuming that, as is commonly the case in practice, the pressure at which the reactants are available is fixed, an increase in the pressure drop across the orifice means can only be achieved by decreasing the pressure within the chamber, which is itself run-desirable both because it leads to an increased tendency towards carbon deposition and because it decreases the rate of reaction. Nevertheless, the disadvantages associated with a reduced pressure will usually be found to be less serious than the disadvantages associated with an unduly low recirculation ratio.

For a given reaction chamber and orifice means, a variation in the velocity of the introduction of the reactants does not have a very marked effect on the recirculation ratio, but an increase in the velocity tends, because of the increase in throughput, to lead to an increase in frictional losses in the chamber and so to a slight decrease in the recirculation ratio. The velocity of introduction of the reactants may be within the range of from 100 feet per second to 2,000 feet per second.

The reaction chamber advantageously comprises a generally cylindrical thermally insulated vessel having mounted coaxially within it a hollow cylindrical member which is shorter than the internal length of the vessel and which divides the interior of the reaction chamber into an inner region of circular cross-section and an outer region of annular cross-section, the two regions being in communication with each other beyond the ends of the hollow cylindrical member, and orifice means arranged to introduce the reactants at or close to one end of the vessel and axially towards the other end of the vessel. With this form of reaction chamber, the rapidly moving stream of reactants entering the reaction chamber passes along the inner region of the chamber, carrying gas already in the reaction chamber along with it, and the moving body of gas then returns along the outer region of the chamber to the vicinity of the orificemeans where it receives fresh impetus from the stream of reactants leaving the orifice means and starts a further cycle of the circulatory movement. Advantageously in order to minimize the resistance afforded to the circulatory movement of the gases by the reaction chamber, the cross-sectional areas (taken in a direction perpendicular to the axis of the reaction chamber) of the two regions of the reaction chamber are substantially equal to one another, or the cross-sectional area of the outer region is slightly greater than that of the inner region.

The orifice means may be situated in the wall of the reaction chamber, but is advantageously formed at the end of a tube or tubes extending within the reaction chamber and parallel to the axis thereof.

The gaseous products are advantageously withdrawn from the reaction chamber at a point that is remote, in a downstream sense, from the orifice means. Thus, when the reaction chamber comprises a generally cylindrical vessel having mounted coaxially within it a hollow cylindrical member, the gaseous products, may be Withdrawn from the reaction chamber through an outlet, situated in or close to the end of the reaction chamber nearest to the orifice means. If desired, a baflle (for example, an annular bafile) may be interposed between the outlet and the adjacent end of the outer region of the reaction chamber in order to lessen any tendency that there may be for gas to enter the outlet before it has completed a single cycle of circulatory movement. A similar result may be achieved by situating the outlet within the outer region of the reaction chamber and facing in the direction of flow of gas in that region. When the orifice means is situated within the inner region of the reaction chamber, that is to say, between the ends of the hollow cylindrical member, the outlet may also be situated within the inner region of the reaction chamber at a point that is a short distance upstream from the orifice means (and is therefore remote, in a downstream sense, from the orifice means).

A reactor suitable for carrying out the process of the invention will now be described by way of example in greater detail with reference to the accompanying drawing which is a diagrammatic axial section of the reactor.

Referring to the drawing, the reactor comprises a vertically mounted cylindrical pressure vessel, which is made up of an open-ended cylindrical steel shell 1 closed by end-pieces 2 and 3 and within which is mounted, coaxially with the shell 1, a hollow heat-resisting steel cylinder 4. Mounted coaxially within the cylinder 4 is a heat-resisting steel cylinder 5.

The cylinder 4 forms a lining for the reaction chamber and the space between the cylinder 4 and the shell 1 contains a thermally insulating material 6. Fixed to the underside of the upper end-piece 2 is a plug 7 of refractory concrete of which the diameter is just less than the internal diameter of the shell 1 and which serves to provide thermal insulation at the top of the reactor. The underside of the plug 7 is recessed over a central circular region of which the diameter is just greater than the external diameter of the cylinder 4 and the outer part of the plug 7 extends a short distance within the upper part of the space between the cylinder and the shell to form a seal with the cylinder 4. Fixed centrally to the top of the lower end-piece 3 is a cylindrical plug 8 of refractory concrete, the diameter of which slightly exceeds the external diameter of the cylinder 4. The plug 8 serves both as a support for the cylinder 4 and to provide thermal insulation at the bottom of the reactor.

At its upper end, the inner tube is secured to the cylinder 4 by a plurality of circumferentially spaced lugs 9 and, at its lower end, the inner tube 5 is located by means of a plurality of circumferentially spaced fins 10 which are secured to the inner tube 5 and bear against the inner surface of the cylinder 4.

The inner tube 5 is shorter than the cylinder 4 and is centrally mounted with respect to the cylinder 4 so that the separation, in an axial direction, between the adjacent ends of these tubes is the same at the top and bottom of the tubes.

The cylinder 4 and the

plugs

7 and 8 together define a reaction chamber, which is divided into an inner region and an outer region by the inner tube 5. The inner and outer regions communicate with one another beyond the ends of the inner tube 5.

An inlet tube 11, which is coaxial with the shell 1, extends through the upper end-piece 2 and the plug 7 and terminates in a nozzle 12. An outlet tube 13 extends vertically through the upper end-piece 2 and the plug 7 and terminates with an open end just above the upper end of the inner tube 5. The outlet tube 13 is so situated that its axis would, if produced, enter the outer region of the reaction chamber.

As an example of suitable dimensions for the reactor shown in the drawing, the shell 1 may be twenty feet in length and approximately thirty inches in diameter. The cylinder 4 may be eighteen feet six inches in length and twelve inches in diameter, and the inner tube 5 may be sixteen feet six inches in length and eight inches in diameter. The internal diameter of the nozzle 12 may be 2 inch.

The following example illustrates the invention.

Using the reactor shown in the drawing and having the dimensions stated to be suitable, a hot hydrogenating gas, that had been obtained by the reaction of butane with steam followed by the replacement of carbon monoxide ,by hydrogen and the removal of carbon dioxide, and consisted of 0.5% carbon dioxide, 1.4% carbon monoxide, 97.3% hydrogen, and 0.8% methane (the percentages being by volume), was introduced into the reaction chamber through the inlet pipe 11 and nozzle 12 until the temperature of the reaction chamber (as measured by a thermocouple situated within the chamber at a point remote from the nozzle 12) reached 600 C. The pressure within the chamber was maintained at 25 atmospheres gauge by a pressure-responsive valve connected to the outlet tube 13. A small proportion of vapour of a light petroleum distillate (having a boiling range of 32 C. to 170 C.) was then incorporated with the hydrogenating gas being supplied to the reaction chamber and the temperature within the reaction vessel steadily increased until i it reached 750 C., which was the desired reaction temperature. Thereupon, the relative proportion of light petroleum distillate vapour in the gases fed to the nozzle 12 was increased gradually to the desired supply rate while the temperature of the entering mixture of hydrogenating gas and oil vapour was gradually reduced in order to maintain the reaction temperature, as measured by thermocouples situated at different points in the reaction chamber, at 750 C.

The final rate of introduction of hydrogenating gas was 7,500 cubic feet per hour (measured at a pressure of 1 atmosphere absolute and a temperature of 15.6 C.) and the final rate of introduction of light petroleum distillate vapour in admixture with the hydrogenating gas was 30 gallons per hour (measured as the liquid petroleum distillate), that is to say, 4 gallons per 1000 cubic feet of hydrogenating gas. The final temperature at which the reactants were introduced was 440 C.

The velocity at which the reactants issued from the nozzle 12 was 1500 feet per second.

The thermocouples were arranged to record the temperature at dififerent points in the interior of the tube 5 (not closer than 3 feet from the nozzle 12) and in the outer region of the reaction chamber, and at no time during hours of continuous operation did the readings of the various thermocouples differ from one another by more than 5 C.

Observations were made on a model into which air was introduced through the nozzle, using a pitot tube to determine the rate of flow of air up the outer annular region at various distances from the axis of the chamber, and thus to determine the total rate of fiow of air up the outer annular region. The observations indicated that, on average, gas introduced into the reaction chamber through the nozzle 12 as reactants recycled within the reaction chamber approximately 20 times before leaving the reaction chamber as products through the outlet 13, that is to say, that the recirculation ratio (as hereinbefore defined) was approximately 20: 1. It should be noted that half the square root of the ratio of the internal crosssectional area of the inner tube to the internal crosssectional area of the nozzle 12 is about 25, thus confirming that this figure is very roughly equal to the recirculation ratio in this case.

The products of reaction on leaving the reaction chamber were cooled and provided 8600 cubic feet per hour of a non-condensable gaseous fraction and 2.7 pounds of condensate for each 1000 cubic feet of the gaseous fraction (the volumes being related to standard temperature and pressure).

On analysis, the gaseous fraction was found to contain 0.4% of carbon dioxide, 1.2% of carbon monoxide, 1.3% of unsaturated hydrocarbons, 35.0% of hydrogen, 49.3% of methane and 12.8% of ethane (the percentages being by volume).

Analysis of the condensate showed that it consisted V of 2.3 pounds of benzene, 0.1 pound of toluene, 0.2

pound of napthalene, and 0.1 pound of higher aromatic hydrocarbons for each 1000 cubic feet of the gaseous fraction.

No deposition of carbon on the walls of the reaction chamber could be detected at the end of the experiment.

What is claimed is:

1. A continuous process for the conversion of the aliphatic constituents of a hydrocarbon distillate oil into a gas containing saturated hydrocarbon constituents selected from the group consisting of methane and mixtures of methane and ethane, which comprises: reacting hydrogen at a temperature ranging from 600 C. to 800 C., under a pressure of at least 20 atmospheres gauge in a thermally insulated reaction chamber with the vapors of said hydrocarbon distillate oil, said chamber defining an endless path along which gas can circulate, the proportion of hydrogen being at least sufficient to substantially completely convert the whole of said aliphatic constituents into said saturated gaseous hydrocarbon constituents, continuously introducing the reactant vapors and gases into an open ended passage within said reaction chamber, said reactants being introduced in the form of a jet through at least one orifice means, to cause a substantial body of gas comprising reactants and reaction products to circulate continuously around said endless path within said reaction chamber due to' the transfer of momentum from the stream of jetted reactants to said body of gas in said chambenpreheating said reactants sufficiently to maintain within said reaction chamber, except in the immediate vicinity of said orifice means, a temperature that is within said temperature range, and that is substantially uniform throughout the interior of said reaction chamber, the overall reaction of said conversion being exothermic; and continuously withdrawing from said reaction chamber, the gaseous products of said reaction.

2. A process as claimed in claim 1, wherein the reaction temperature is within the range of from 700 C. to 800 C.

3. A process as claimed in claim 1, wherein the gas comprising hydrogen consists substantially wholly of hydrogen.

4. A process as claimed in claim 1, wherein the gas comprising hydrogen is a gaseous mixture consisting mainly of hydrogen.

5. A process as claimed in claim 1, wherein the gas comprising hydrogen also contains steam.

6. A process as claimed in claim 1, wherein the oil vapour is that of a light petroleum distillate, and the distillate vapour introduced with each 1000 cubic feet, measured at S.T.P., of hydrogen introduced in the gas comprising hydrogen is within the range of from 2 to 8 gallons, measured as liquid distillate.

7. A process as claimed in claim 6, wherein the gas comprising hydrogen also contains steam.

8. A process as claimed in claim 1, wherein the mixed preheat temperature of the oil vapour and the gas comprising hydrogen is at least 350 C. and below the reaction temperature.

9. A process as claimed in claim 8, wherein the said mixed preheat temperature is within the range of from 400 C. to 550 C.

10. A process as claimed in claim 1, wherein the ratio of the volume of gas circulating within the reaction chamber to the volume of gas withdrawn therefrom during one complete period of the circulatory motion is at least 3:1.

11. A process as claimed in claim 10, wherein the said ratio is at least 1021.

12. A process as claimed in claim 10, wherein the said ratio is at least 20:1.

13. A process as claimed in claim 1, wherein the tem- V perature within the chamber is, except in the immediate vicinity of the orifice means, uniform to within :10 C.

14. A process as claimed in claim 1, wherein the temperature within the chamber is, except in the immediate vicinity of the orifice means, uniform to within i5 C.

15. A process as claimed in claim 1, wherein the tem-.

perature within the chamber is, except in the immediate vicinity of the orifice means, uniform to within :25 C.

16. A process as claimed in claim 1, wherein the ratio of the cross-sectional area of the passage into which the reactants are introduced to the cross-sectional area of the orifice means, each being taken in a plane that is perpendicular to the direction of flow of the gases, is at least :1.

17. A process as claimed in claim 16, wherein the said ratio is at least 400:1.

18. A process as claimed in claim 1, wherein the reactants are introduced into the reaction chamber at a velocity within the range of from 100 feet per second to 2,000 feet per second.

19. A process as claimed in claim 1, wherein the gaseous products are withdrawn from the reaction chamber at a point that is remote, in a downstream sense, from the orifice means. References Cited UNITED STATES'PATENTS 1,961,288 6/1934 Faber 208.133 2,759,806 -8/ 1956 Pettyjohn et al. 48213 2,786,877 3/1957 King 208-106 3,025,149 3/1962 Eastman 48213 3,148,135 9/ 1964 Schlinger et al. 208-107 HERBERT LEVINE, Primary Examiner. DELBERT E. GANTZ, Examiner,