WO2025170892A1

WO2025170892A1 - Composition of matter for gas-stripping oxo aldehydes hydroformylation process

Info

Publication number: WO2025170892A1
Application number: PCT/US2025/014424
Authority: WO
Inventors: Alan George Wonders; Kevin Scott HOWE; Debangshu GUHA; Scott Allen OWENS; Kevin John FONTENOT; Kenny Randolph Parker; Jody Lee Rodgers; Beth Williford ALDERSON
Original assignee: Eastman Chemical Co
Current assignee: Eastman Chemical Co
Priority date: 2024-02-06
Filing date: 2025-02-04
Publication date: 2025-08-14
Anticipated expiration: 2026-08-06

Abstract

A hydroformylation composition produced by the process is provided comprising forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein the multiphase reaction medium comprises one or more olefins, molecular hydrogen, carbon monoxide and a catalyst, wherein the hydroformylation involves the net addition of a formyl group and a hydrogen to a carbon-carbon double bond to form two aldehyde products. Further, this invention provides for production of oxo aldehydes, especially propionaldehyde and butyraldehyde, in a hydroformylation reaction medium and vessel using an improved process and apparatus.

Description

COMPOSITION OF MATTER FOR GAS-STRIPPING OXO ALDEHYDES HYDROFORMYLATION PROCESS BACKGROUND OF THE INVENTION [0001] This invention provides for production of oxo aldehydes, especially propionaldehyde (HPr) and butyraldehyde (HBu), in a hydroformylation reaction medium and vessel using an improved process and apparatus. Hydroformylation involves the net addition of a formyl group (CHO) and a hydrogen to a carbon-carbon double bond to form two aldehyde products. Aldehydes are used in various end uses such as, for example, production of plastics, alkyd resins, rubber chemicals, disinfectants, preservatives, alcohols, acids, plasticizers, fragrances, drugs, and pesticides. There is a need in the industry for more efficient processes to produce oxo aldehydes. [0002] Processes are disclosed that utilize a high gas flow rate though a hydroformylation reaction medium and vessel resulting in one or more of the following improvements while still maintaining necessary gas-only stripping rates of aldehyde: improved reaction zone temperatures, improved reaction zone temperature homogeneity, simplification of the entire process as compared to known processes, reduced by-product formation, reduce ligand degradation, and improved the n/i ratio of HBu. In one embodiment, these improvements can be implemented in a Bubble Column Reactor (BCR). The BCR can have a relatively short height (H) and small H/D to limit vessel cost, compressor power for mixing, mass of liquid phase of reaction medium, and axial thermal gradients therein. [0003] A hydroformylation reaction is exothermic and therefore requires continuous heat removal from the multi-phase reaction medium. Improved reaction zone temperature homogeneity is desired to minimize unwanted chemical reactions including alkene hydrogenation to alkane; ligand degradation, and Rh-ligand purge loss. Furthermore, minimizing thermal gradients within a reaction medium enables better optimization of hydroformylation reaction rates versus adverse chemistry reaction rates. [0004] One advantage of this invention is heat removal from the reaction zone is accomplished in a fashion that minimizes thermal gradients within the reaction zone. In one embodiment, enough gas flow is provided to facilitate removal of virtually the entire reaction zone heat duty into a recycle gas cooling condenser. In another embodiment, at least a portion of the reaction zone heat duty is removed by routing multi-phase reaction medium from the reaction zone to an external heat exchanger. [0005] These oxo processes provide distinctly greater gas flow rates through a reaction medium and vessel per unit production of aldehyde (HBu, HPr) than is presently practiced. This is surprising and counterintuitive to normal chemical engineering for minimizing cost because: 1) greater gas flow rates through a reaction medium result in higher recycle gas flow rates which increase the capital and operating costs for the gas recycle compressor, and 2) reduction in the conversion per pass of reactant gases can possibly lead to an increased concentration and yield loss of at least one of the valuable reactant gases via the purge gas flow used to limit the accumulation of non- reactive inerts in the recycle gas. Although this adds compressor capital cost and compression horsepower, it is surprisingly cheaper on both capital and utilities than adding a separate liquid stripping system for recovering aldehyde product when using cooler reaction temperatures and conventional, smaller gas recycle rates. [0006] In addition to cooling a reaction medium, the much higher gas flow rate markedly reduces the conversion per pass of feed gases CO/H2/olefin, making inlet and outlet compositions closer together. Although this is instinctively adverse for minimizing purge gas yield loss, it is nonetheless often surprisingly important for optimization of reaction chemistry more near the ideal values for greatest n/i ratio and greatest hydroformylation reaction intensity. This can be particularly important for CO when using some of the Rh-ligand pairs that provide production of very high n/i product ratios, because a rather narrow range of CO partial pressures is often needed to balance inter-phase mass transfer rates with the liquid phase concentration of dissolved CO that are needed for greatest reactivity and largest n/i ratio. [0007] The reduction in conversion per pass of olefin can be adverse for olefin yield, since the olefin concentration in reaction gas outlet, which is source of purge gas, will be closer to the inlet concentration. However, maximizing the olefin conversion and yield can often be more efficiently handled, even increased, by provision of a separate, relatively small, and very highly staged purge gas scrubber reactor, particularly a tall H/D tray staged BCR. In one embodiment, a high gas flow hydroformylation reactor is coupled to a purge gas scrubber reaction medium and vessel that uses oxo hydroformylation to convert reactant gases purged from a first, main oxo reaction medium and vessel into additional aldehyde product. BRIEF DESCRIPTION OF THE DRAWING(S) [0008] Figure 1 shows one embodiment of a purge bubble column reactor. [0009] Figure 2 shows one embodiment of a purge reactor and from multiple main reactors. [0010] Figure 3 shows another embodiment of a purge reactor. [0011] Figure 4 shows one embodiment of purge reactor and first, main reactor. [0012] Figure 5 shows one embodiment of the main reactor and purge reactor for an oxo process. DETAILED DESCRIPTION OF THE INVENTION [0013] “Hydroformylation” is defined herein as combination by chemical reaction of at least one olefin plus carbon monoxide (CO) and hydrogen (H2) to form at least one aldehyde product wherein said chemical reaction employs at least one catalyst compound comprising at least one transition metal. The aldehyde product can at least partially be comprised in a condensed liquid phase. [0014] “Multi-phase hydroformylation reaction medium”, “aerated hydroformylation reaction medium”, “hydroformylation reaction medium”, and more succinctly “reaction medium” are defined herein as equivalent terms meaning any reaction medium comprising a gaseous phase and at least 1 liquid phase. Owing to the gaseous tendency of CO and H2, a hydroformylation reaction medium is typically at least a 2-phase reaction medium comprising a liquid phase and a gaseous phase wherein the gaseous phase comprises gaseous CO and H2. In addition, hydroformylation reaction medium may sometimes comprise multiple liquid phases, e.g. a water-rich phase and an aldehyde-rich phase, and sometimes even a solid phase, e.g. a supported solid-phase catalyst or a particulate used to control the gas hold-up within reaction medium or to reduce foaming atop reaction medium. [0015] A “hydroformylation reaction means”, “hydroformylation reactor”, and more succinctly “reactor” are defined herein as equivalent terms meaning any vessel, conduit, or other structure containing a continuous portion of reaction medium. Suitable reactors include, for example, bubble-agitated reactors (e.g., bubble column reactors), mechanically agitated reactors (e.g., continuous stirred tank reactors), and flow agitated reactors (e.g., jet reactors and static mixer reactors). [0016] A “hydroformylation reactor process” is defined herein as the combination of at least one reactor in combination with associated reaction feeds supplying at least one olefin plus CO and H2; at least one apparatus for removing heat, for the hydroformylation reaction is exothermic; at least one source providing at least one transition metal hydroformylation catalyst into reaction medium; and at least one means of separating and recovering a liquid- phase crude aldehyde product substantially free of catalyst. Many times, but not necessarily, a hydroformylation reactor process will also have associated recycle flows of various gaseous and/or liquid streams that have been removed from a reactor, processed in various ways to provide at least a partial separation of compounds, and with portions thereof being returned to the reactor and its contained reaction medium. [0017] As used and defined herein, the adjectives “first” and “second” modifying reactors, reaction media and reactor process refer to the relative sequence of the two subjects and do not preclude that there may be additional upstream and downstream reactors, reaction media and reactor processes. [0018] In particular, a “second reactor”, “second reaction medium” or “second reactor process” are defined herein as comprising conversion to aldehyde product of at least a portion of a feed gas mixture comprising olefin, CO, and H2 that is essentially derived from olefin, CO and H2 that have already passed unreacted through at least one first reactor, reaction medium, and reactor process, respectively. In addition, the words “scrubbing” and “polishing” are defined as equivalent to “second” when used in this context of reactor, reaction medium and reactor process. [0019] “Gaseous recycle mixture” and “recycle gas” are defined herein as equivalent terms meaning a gaseous mixture essentially obtained from a portion of reaction medium and essentially having been formed outside said reaction medium and wherein at least a portion of said gaseous mixture is then returned to said reaction medium. Examples comprise obtaining a gaseous mixture by separation from liquid portions of reaction medium in a gravitational disengaging zone followed by compressing at least a portion of said gaseous mixture and returning at least a portion thereof to its source reaction medium; obtaining a gaseous mixture by removing at least a portion of a liquid phase of reaction medium outside of a reactor, reducing the pressure of said liquid portion to obtain a portion of gaseous mixture comprising evaporated liquid compounds, separating said gaseous mixture from remaining liquid portion in a gravitational disengaging zone followed by compressing at least a portion of said gaseous mixture and returning at least a portion thereof to its source reaction medium; and so on. [0020] “Reactor purge gas” and “purge gas” are defined herein as equivalent terms meaning a portion of gaseous mixture essentially derived from a portion reaction medium that is not thereafter returned to said portion of reaction medium. Quite often purge gas is derived by dividing away a smaller portion of a stream mainly comprising recycle gas. The process function of purge gas is often to remove an accumulation of inert gases within recycle gas; but sometimes purge gas serves other purposes, e.g., adjusting the portions of olefin, CO, and H2 within recycle gas, adjusting the heat capacity of recycle gas, and so on. [0021] As known in the art and as defined herein, "gas hold-up" is the volume fraction of a multi-phase medium that is in the gaseous state. [0022] As known in the art and as defined herein, “synthesis gas” is a gaseous mixture comprising at least 90 mol% CO and H2, although the ratio of CO:H2 is highly variable depending upon the source and the intended usage of a particular portion of synthesis gas. [0023] Aldehyde formed is defined herein as the full amount of aldehyde created by carbonylation reactions combining equimolar amounts of at least one olefin, CO, and H2 and without reduction for any of this aldehyde that may subsequently react to form a byproduct other than aldehyde, said byproducts comprising alcohols, acids, and heavy byproducts from aldehyde condensation reactions. [0024] The aldehyde formation heat duty is defined herein as the enthalpy difference between equimolar amounts of gaseous CO, H2, and at least one olefin compared with the resulting equimolar aldehyde vapor, all evaluated at the temperature and pressure of the highest elevation of said reaction medium, and notwithstanding that hydroformylation reaction typically occurs in the liquid phase. [0025] The “catalyst metal” is selected from the Group VIII transition metals and may be provided in the form of various metal compounds such as carboxylate salts of the transition metal. Rhodium is the preferred Group VIII metal. The source of rhodium for the active catalyst include rhodium(II) or rhodium(III) salts of carboxylic acids, examples of which include di-rhodium tetraacetate dihydrate, rhodium(II) acetate, rhodium(II) isobutyrate, rhodium(II) 2-ethylhexanoate, rhodium(II) benzoate and rhodium(II) octanoate. Also, rhodium carbonyl species such as Rh4 (CO)12, Rh6 (CO)16 and rhodium(I) acetylacetonate dicarbonyl may be suitable sources of rhodium. Additionally, rhodium organophosphine complexes such as tris(triphenylphosphine) rhodium carbonyl hydride may be used when the phosphine moieties of the complex feed are easily displaced. Other rhodium sources include rhodium salts of strong mineral acids such as chlorides, bromides, nitrates, sulfates, phosphates and the like. Rhodium 2-ethylhexanoate is a particularly preferred source of rhodium from which to prepare the complex catalyst of the invention because it is a convenient source of soluble rhodium, as it can be efficiently prepared from inorganic rhodium salts such as rhodium halides. [0026] The “catalyst ligand” is defined as an organophosphorus molecule that is capable of coordinating to the catalyst metal. Catalyst ligands and their methods of preparation are well known in the art with the most common examples being organophosphines (monomeric as well as chelating forms), organophosphites (monomeric as well as chelating forms), halophosphites and phosphoramidites. [0027] The “catalyst compound” is defined as the combination of catalyst metal and catalyst ligand to form the active hydroformylation catalyst. In general, hydroformylation reactions involve the production of aldehydes by reacting an olefinic unsaturated compound with carbon monoxide and hydrogen in the presence of a catalyst compound homogeneously dissolved in a liquid medium that also contains a solvent for the catalyst, and free catalyst ligand (i.e., ligand that is not complexed with the rhodium metal in the active complex catalyst) that is likewise homogeneously dissolved. The hydroformylation process encompassed by this invention may be carried out in any excess amount of free catalyst ligand desired with at least one mole of free catalyst ligand per mole catalyst metal present in the reaction medium. In general amounts of catalyst ligand of from about 4 to about 50 moles per mole rhodium present in the reaction medium should be suitable for most purposes, said amounts being the sum of both the amount of catalyst ligand that is bound (complexed) to the rhodium present and the amount of free (non-complexed) catalyst ligand present. Of course, if desired, make-up catalyst ligand can be supplied to the reaction medium of the hydroformylation process, at any time and in any suitable manner, to maintain a predetermined level of free catalyst ligand in the reaction medium. Moreover, it is to be understood that while the catalyst ligand of the catalyst compound and excess free catalyst ligand in a given process are both normally the same, different catalyst ligands, as well as, mixtures of two or more different catalyst ligands may be employed for each purpose in any given process, if desired. [0028] “Heavy condensation product of aldehydes” comprise the higher boiling by-products that are naturally formed during the process of the hydroformylation reaction and the subsequent steps, e.g., distillations, that are required for aldehyde product isolation such as alcohols, esters, acetals and hydroxyaldehydes which are retained as high boiling liquids in the hydroformylation reactor or at the bottom of subsequent distillation columns. These products derive ultimately from the product aldehyde via a variety of reactions (e.g., aldol, Tishchenko, hydrogenation, acetalization, Cannizzaro) but do not include the catalyst compound or any of its decomposition products. [0029] “Catalyst decomposition product” is defined as any by-product that derives from the catalyst ligand (i.e., contains molecular fragments originating from the catalyst ligand). These decomposition products may be formed by a variety of chemistries with the most common being (1.) oxidation of trivalent organophosphorus molecules to the corresponding pentavalent oxide in the presence of oxidants such as oxygen or peroxides and (2.) cleavage of phosphorus oxygen bonds by hydroxyl-containing molecules (e.g., water, alcohols, or hydroxyl-containing heavy condensation products of aldehydes). Theses decomposition agents can be formed during hydroformylation or enter via contamination of the reactor feeds (e.g., olefin or syngas). Oxo Reactor Gas Flow Ratio to Aldehyde Production [0030] In one embodiment of the invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein the multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein the maximum inside diameter of the reaction medium is greater than 1.5 meters; (b) reacting at least a portion of the one or more olefins, the H2, and the CO in the multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; and (c) withdrawing from the hydroformylation reactor at least a portion of gas phase effluent comprising at least a portion of the one or more aldehydes and at least one non-aldehyde component, wherein the ratio of molar flow rate of the non-aldehyde gas phase effluent components withdrawn in the step (c) to the molar formation rate of the one or more aldehydes formed in the step (b) is at least 18:1 and not more than 36:1; (d) cooling at least a portion of the gas phase effluent to form at least a portion of cooled gas phase effluent and at least a portion of condensed crude aldehyde liquid phase comprising at least a portion of at least one condensed aldehyde and at least a portion of at least one olefin dissolved therein; (e) separating and compressing at least a portion of the cooled gas phase effluent to form a gas phase recycle gas; (f) returning at least a portion of the recycle gas of step (e) to the multiphase reaction medium of step (b). The molar ratio of the condensed aldehyde in step (d) can be at least 1.06, 1.07, 1.08, 1.09, or 2 times the amount of the formed aldehyde in step (b). Also, the compression of the step (e) can provide a pressure increase of at least 80 kPa and less than 450 kPa. In addition, the compression of step (e) can provide an amount of work per kg-mole of the formed aldehyde of step (b) that is at least 0.6 kilowatts per kg-mole and less than 1.7 kilowatts per kg-mole. [0031] In another embodiment of the invention, the liquid phase of step (a) comprises at least 45 and less than 85 mole percent aldehyde. [0032] In another embodiment, propylene comprises the preponderance of the one or more olefins. The ratio of molar flow rate of the non-aldehyde gas phase effluent components to the molar formation rate of the one or more aldehydes can be from at least 20:1 and not more than 28:1. In another embodiment, the ratio of molar flow rate of the non-aldehyde gas phase effluent components to the molar formation rate of the one or more aldehydes is at least 21:1 and not more than 28:1. [0033] In one embodiment of the invention, the hydroformylation reaction medium is comprised within a bubble column reactor. The amount of aldehyde formed in the step (b) can be produced at a rate of at least 25 and less than 100 kg-moles per hour per square meter of cross-sectional area, wherein the cross-sectional area is calculated in as Pi/4 * D^2 using the maximum diameter D. In addition, the molar flow rate of aldehyde in the gas phase effluent of step (b) can range from 1.1 to 1.7 times, from 1.1 to 1.69 times, from1.1 to 6 times, from 1.1 to 5 times, 1.1 to 4, 1.1 to 3 times, or 1.1 to 2 times the molar formation rate of aldehyde in the step (b). Oxo Reactor Gas Flow Ratio to Propionaldehyde Production [0034] A hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein the multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein the maximum inside diameter of the reaction medium is greater than 1.5 meters; (b) reacting at least a portion of the one or more olefins, the H2, and the CO in the multiphase reaction medium to thereby form one or more aldehydes comprising predominantly propionaldehyde via hydroformylation; and (c) withdrawing from the hydroformylation reactor at least a portion of gas phase effluent comprising at least a portion of the propionaldehyde and at least one non-aldehyde component, wherein the ratio of molar flow rate of the non-aldehyde gas phase effluent components withdrawn in the step (c) to the molar formation rate of the one or more aldehydes formed in the step (b) is at least 7:1 and not more than 22:1,can be at least 8:1 and not more than 20:1, at least 9:1 and not more than 18:1, and at least 10:1 and not more than 16:1. [0035] In the hydroformylation process, the reaction medium can be comprised within a bubble column reactor. [0036] The aldehyde formed in the step (b) can be produced at a rate of at least 25 and less than 100 kg-moles per hour per square meter of cross- sectional area, wherein the cross-sectional area is calculated in as Pi/4 * D^2 using the maximum diameter D. Also, the molar flow rate of aldehyde in the gas phase effluent of step (b) can be at least 1.1 times and less than 1.7 times the molar formation rate of aldehyde in the step (b). The molar ratio of the condensed aldehyde in step (d) can be at least 1.06 times the amount of the formed aldehyde in step (b). In addition, the compression of the step (e) provides a pressure increase of at least 80 kPa and less than 450 kPa. The compression of step (e) can provide an amount of work per kg-mole of the formed aldehyde of step (b) that is at least 0.6 kilowatts per kg-mole and less than 1.7 kilowatts per kg-mole. The liquid phase of step (a) can comprise at least 30 and less than 75 mole percent propionaldehyde. Oxo Reactor Gas Velocity [0037] In one embodiment of the invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein the reactor has a maximum diameter D that is at least 1.5 meters and comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (b) reacting at least a portion of the one or more olefins, the H2, and the CO in the multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; (c) withdrawing from the hydroformylation reactor at least a portion of a gas phase effluent comprising at least a portion of the one or more aldehydes and non-aldehyde components, wherein the volumetric flow (m3/s) of all gas phase effluents withdrawn from the reactor divided by the horizontal cross-sectional area (m2) of the reactor evaluated at the maximum diameter D of step (a) is at least 0.18 and not more than 0.8 meters per second (m/s); [0038] at least 0.21 and not more than 0.7 meters per second (m/s), at least 0.24 and not more than 0.6 meters per second (m/s), and at least 0.27 and not more than 0.6 meters per second (m/s). (d) cooling at least a portion of the gas phase effluent to form at least a portion of cooled gas phase effluent and at least a portion of condensed crude aldehyde liquid phase comprising at least a portion of at least one aldehyde and at least a portion of at least one olefin dissolved therein; (e) separating and compressing at least a portion of the cooled gas phase effluent to form a gas phase recycle gas; (f) returning at least a portion of the recycle gas of step (e) to the multiphase reaction medium of step (b). [0039] In this process, the reaction medium can be comprised within a bubble column reactor and the aldehyde formed comprises predominately propionaldehyde and butyraldehyde. In other embodiments of the invention, the hydroformylation reactor can have a maximum diameter D in millimeters of at least 6.0 times and less than 10.0 times the kg-moles per hour of aldehyde formed in the step (b). The gas agitation power in the reactor can be at least 0.3 and not more than 6.0 kilowatts per cubic meter of the liquid phase of multiphase reaction medium. The gram-mole per hour of the aldehyde formed in step (b) divided by the liters of the liquid phase of multiphase reaction medium in step (a) can be at least 7 and is less than 26 (g-mole aldehyde formed/hr/L liquid phase). [0040] The compression in step (e) can provide a pressure increase of at least 80 kPa and less than 450 kPa. In addition, the ratio of molar flow rate of the non-aldehyde gas phase effluent components withdrawn in the step (c) to the molar formation rate of the one or more aldehydes formed in the step (b) is at least 18:1 and not more than 36:1. The gram-mole per hour of the aldehyde formed in step (b) divided by the liters of the liquid phase of multiphase reaction medium in step (a) is at least 8 and can be less than 24 (g-mole aldehyde formed/hr/L liquid phase). Oxo Thermal Duty Using Large Gas Flows [0041] In another embodiment of the invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein the multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein the maximum diameter of the multiphase reaction medium is at least 1.5 meters; (b) reacting at least a portion of the one or more olefins, the H2, and the CO in the multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; (c) withdrawing from the hydroformylation reactor at least a portion of a gas phase effluent comprising at least a portion of the one or more aldehydes and at least a portion of non-aldehyde components; (d) cooling at least a portion of the gas phase effluent, wherein the enthalpy exchanged out of the gas phase effluent is at least 15 percent, at least 35 percent, at least 45 percent, at least 50 percent, at least 65 percent, at least 90 and not more than 125 percent, at least 98 and not more than 120 percent , at least 102 and not more than 115 percent, or at least 104 and not more than 110 percent of the enthalpy of reaction provided in the reacting step (b). [0042] This process can additionally comprise further steps. In one embodiment, the process further comprising a step (e) compressing at least a portion of the cooled gas phase effluent. In another embodiment, the process further comprises a step (f) returning at least a portion of the cooled and compressed gas phase effluent to the multiphase reaction medium of step (b). [0043] In this hydroformylation process during the cooling step (d), at least a portion of the enthalpy exchanged out of the gas phase effluent is exchanged into at least a portion of other fluid via a solid cooling surface in contact with the gas phase effluent and the other fluid. [0044] In embodiments of this process, the aldehyde formed comprises predominately propionaldehyde and butyraldehyde. In addition, the reaction medium of step (b) is comprised within a bubble column reactor. Oxo Reactor Gas Compositions For Larger Gas Flow [0045] In another embodiment, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein the maximum diameter of said multiphase reaction medium is at least 1.5 meters, wherein the time-averaged superficial velocity of said gas phase at the half-height within said multiphase reaction medium is at least 0.13 and not more than 0.7 meters per second (m/s); (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; (c) withdrawing from said hydroformylation reactor at least a portion of a gas phase effluent comprising at least a portion of said one or more aldehydes and at least a portion of non-aldehyde components, wherein the ratio of the mole fraction sum of said non-aldehyde gas phase effluent components divided by the mole fraction sum of said aldehyde gas phase effluent components is at least 3:1 and not more than 32:1; (d) cooling at least a portion of said gas phase effluent, wherein the enthalpy exchanged out of said gas phase effluent is at least 35 percent of the enthalpy of reaction provided in said reacting step (b). [0046] In these hydroformylation processes having large gas flows during said cooling step (d), at least a portion of the enthalpy exchanged out of said gas phase effluent is exchanged into at least a portion of other fluid using a solid cooling surface in contact with said gas phase effluent and said other fluid. [0047] These hydroformylation processes can additionally comprise a step (e) compressing at least a portion of said cooled gas phase effluent. [0048] In yet another embodiment of these inventions, the hydroformylation processes additionally comprise a step (f) returning at least a portion of said cooled and compressed gas phase effluent to said multiphase reaction medium of step (b). [0049] The aldehyde formed in these processes can comprise predominately propionaldehyde and butyraldehyde. [0050] In another embodiments of these inventions, in step (c), the ratio of the mole fraction sum of said non-aldehyde gas phase effluent components divided by the mole fraction sum of said aldehyde gas phase effluent components is at least 6:1 and not more than 28:1. Other ranges for the mole fraction sum of said aldehyde gas phase effluent components is at least 9:1 and not more than 24:1 and at least 12:1 and not more than 20:1. [0051] The reaction medium of step (b) can be comprised within a bubble column reactor. A bubble column reactor is disclosed in U.S. Patent Number 7,910,071, herein incorporated by reference. [0052] In this process, during the cooling step (d), at least a portion of the enthalpy exchanged out of said gas phase effluent is exchanged into at least a portion of other fluid via a solid cooling surface in contact with said gas phase effluent and said other fluid. [0053] In this embodiment, the process can additionally comprise a step (e) compressing at least a portion of said cooled gas phase effluent. The process can also additionally comprising a step (f) returning at least a portion of said cooled and compressed gas phase effluent to said multiphase reaction medium of step (b). Oxo Reactor Gas Agitation Power [0054] A hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation to produce a gas phase effluent; wherein the gas agitation power is at least 0.2 and not more than 6.0 kilowatts per cubic meter of said reaction medium; and (c) withdrawing from said hydroformylation reactor at least a portion of said gas phase effluent comprising at least a portion of said one or more aldehydes and non-aldehyde components. [0055] In other embodiments, the gas agitation power can be at least 0.2 and not more than 5.5, at least 0.2 to not more 5.0, at least 0.2 to not more than 4.5, at least 0.2 to not more than 4.0, at least 0.2 to not more than 3.5, at least 0.2 to not more than 3.0, at least 0.2 to not more than 2.5, at least 0.3 and not more than 6.0, at least 0.3 and not more than 5.5, at least 0.3 to not more 5.0, at least 0.3 to not more than 4.5, at least 0.3 to not more than 4.0, at least 0.3 and not more than 3.5, at least 0.3 and not more than 3.0, at least 0.3 and not more than 2.5, at least 0.5 and not more than 6.0, at least 0.5 and not more than 5.5, at least 0.5 to not more 5.0, at least 0.5 to not more than 4.5, at least 0.5 to not more than 4.0, at least 0.5 and not more than 3.5, at least 0.5 and not more than 3.0, at least 0.5 and not more than 2.5at least 0.9 and not more than 6.0, at least 0.9 and not more than 5.5, at least 0.9 to not more 5.0, at least 0.9 to not more than 4.5, at least 0.9 to not more than 4.0, at least 0.9 and not more than 3.5, at least 0.9 and not more than 3.0, at least 0.9 and not more than 2.5, at least 1.0 and not more than 6.0, at least 1.0 and not more than 5.5, at least 1.0 to not more 5.0, at least 1.0 to not more than 4.5, at least 1.0 to not more than 4.0, at least 1.0 and not more than 3.5, at least 1.0 and not more than 3.0, at least 1.0 and not more than 2.5, at least 1.4 and not more than 6.0, at least 1.4 and not more than 5.5, at least 1.4 to not more 5.0, at least 1.4 to not more than 4.5, at least 1.4 to not more than 4.0, at least 1.4 and not more than 3.5, at least 1.4 and not more than 3.0, at least 1.4 and not more than 2.5. kilowatts per cubic meter of said reaction medium. [0056] In another embodiment of this invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation to produce a gas phase effluent; wherein the gas agitation power is at least 0.2 and not more than 6.0 kilowatts per cubic meter of said reaction medium; and (c) withdrawing from said hydroformylation reactor at least a portion of said gas phase effluent comprising at least a portion of said one or more aldehydes and non-aldehyde components; (d) cooling at least a portion of said gas phase effluent; (e) compressing at least a portion of said cooled gas phase effluent; (f) returning at least a portion of said cooled and compressed gas phase effluent to said multiphase reaction medium of step (b). [0057] In any of these embodiments, the aldehyde formed comprises predominately propionaldehyde and butyraldehyde. In addition, the reaction medium is comprised within a bubble column reactor. Oxo Reactor Axial Temperature Gradient [0058] In an embodiment of the invention, a hydroformylation process is provided where the reactor has an axial temperature gradient. The process comprises: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein said reaction medium has a highest elevation, a lowest elevation, and a maximum diameter, wherein said diameter is at least 1.5 meters; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; (c) withdrawing from said hydroformylation reactor at least a portion of said gas phase effluent comprising at least a portion of said one or more aldehydes and non-aldehyde components, wherein said gas phase effluent is withdrawn above an elevation that is 0.5 times said maximum diameter below said highest elevation of said multiphase reaction medium; (d) optionally, cooling at least a portion of said gas phase effluent; (e) optionally, compressing at least a portion of said cooled gas phase effluent; (f) optionally, returning at least a portion of said cooled and compressed gas phase effluent to said multiphase reaction medium. (g) determining a time-averaged and area-averaged temperature of said multiphase reaction medium of said step (b) at an elevation that is 0.5 times said maximum diameter above said lowest elevation of said multiphase reaction medium; (h) determining a time-averaged and volume-averaged temperature of said gas phase effluent of said step (c); wherein the difference between said time-averaged and area-averaged temperature determined in said step (g) and said time-averaged and volume-averaged temperature determined in said step (h) is at least 1°C and not more than 16°C. [0059] In other embodiments of this invention, the difference between said time-averaged and area-averaged temperature determined in said step (g) and said time-averaged and volume-averaged temperature determined in said step (h) is at least 1°C and not more than 12°C, at least 1°C and not more than 8°C, or at least 1°C and not more than 4°C . [0060] The aldehyde formed can comprise predominately propionaldehyde and butyraldehyde. In addition, the reaction medium can be comprised within a bubble column reactor. A bubble column reactor was previously described in this disclosure. [0061] In another embodiment of this invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein said reaction medium has a highest elevation, a lowest elevation, and a maximum diameter, wherein said maximum diameter is at least 1.5 meters; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; (c) withdrawing from said hydroformylation reactor at least a portion of said gas phase effluent comprising at least a portion of said one or more aldehydes and non-aldehyde components, wherein said gas phase effluent is withdrawn above an elevation that is 0.25 times said maximum diameter below said highest elevation of said multiphase reaction medium; (d) cooling at least a portion of said gas phase effluent; (e) compressing at least a portion of said cooled gas phase effluent; (f) returning at least a portion of said cooled and compressed gas phase effluent to said multiphase reaction medium. (g) determining a time- averaged and area-averaged temperature of said multiphase reaction medium of said step (b) at an elevation that is 0.25 times said maximum diameter above said lowest elevation of said multiphase reaction medium; (h) determining a time-averaged and volume-averaged temperature of said gas phase effluent of said step (c); wherein the difference between said time-averaged and area- averaged temperature determined in said step (g) and said time-averaged and volume-averaged temperature determined in said step (h) is at least 1°C and not more than 16°C. Oxo Optimized Process [0062] In one embodiment of this invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein the multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (b) reacting at least a portion of the one or more olefins, the H2, and the CO in the multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; (c) withdrawing from the hydroformylation reactor at least a portion of a gas phase effluent comprising at least a portion of the one or more aldehydes and non- aldehyde components; (d) cooling at least a portion of the gas phase effluent to form at least a portion of cooled gas phase effluent and at least a portion of condensed crude aldehyde liquid phase comprising at least a portion of at least one aldehyde and at least a portion of at least one olefin dissolved therein; (e) separating and )compressing at least a portion of the cooled gas phase effluent to form a gas phase recycle gas; (f) returning at least a portion of the recycle gas of step (e) to the multiphase reaction medium of step (b); (g) returning at least a portion of condensed crude aldehyde liquid to the hydroformylation reactor of step (a); (h) contacting in a syngas scrubber at least a portion of a gas phase raw syngas feed comprising H2, CO, and at least one syngas impurity with at least a portion of the condensed crude aldehyde liquid of step (d) to form a liquid phase of improved crude aldehyde and a gas phase of improved syngas feed wherein at least a portion of at least one of the dissolved olefins of the condensed crude aldehyde liquid is transferred into the improved syngas feed; (i) supplying at least a portion of the improved syngas feed from step (g) to the multiphase reaction medium of step (b) or to the recycle gas of step (e); (j) supplying at least a portion of at least one olefin feed supply to at least a portion of the reaction medium of step (b) or to at least a portion of the recycle gas of step (e) or to syngas-recycle gas mix or to improved syngas; (k) optionally, separating at least a portion of cooled gas phase effluent of step (d) or at least a portion of the recycle gas phase effluent of step (e) to form a purge gas. Oxo Reactor Gas Phase Gradients [0063] A hydroformylation process comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein said reaction medium has a highest elevation, a lowest elevation, and a maximum diameter, wherein said maximum diameter is at least 1.5 meters, wherein the greatest partial pressure of said CO in said gas phase present at a location within 0.5 meters of said lowest elevation is not more than 500 kilopascals (kPa); (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation, wherein the time-averaged superficial velocity of said gas phase at the half-height within said multiphase reaction medium is at least 0.07 and not more than 0.8 meters per second (m/s); (c) withdrawing from said hydroformylation reactor at least a portion of said gas phase comprising at least a portion of said one or more aldehydes and non- aldehyde components, wherein said gas phase effluent is withdrawn at an elevation that is not more than 2 meters below said highest elevation of said multiphase reaction medium; wherein the partial pressure of CO in said portion of effluent gas phase from said step (c) divided by said greatest partial pressure of CO from said step (a) is at least 0.30 and not more than 0.85. [0064] A hydroformylation process comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein said reaction medium has a maximum diameter that is at least 1.5 meters, wherein the greatest partial pressure of said CO in said gas phase present at a location within 0.5 meters of said lowest elevation is not more than than 500 kilopascals (kPa); (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; (c) withdrawing from said hydroformylation reactor at least a portion of said gas phase comprising at least a portion of said one or more aldehydes and non-aldehyde components, wherein the partial pressure of said CO in said gas phase is not more than 400 kilopascals (kPa); (d) optionally, cooling at least a portion of said gas phase effluent; (e) optionally, compressing at least a portion of said cooled gas phase effluent; (f) optionally, returning at least a portion of said cooled and compressed gas phase effluent to said multiphase reaction medium; (g) feeding at least one portion of a gas phase comprising CO into said multiphase reaction medium of step (b); (h) determining the total molar flow rate of all gas phase CO fed into the reaction medium of step (b); (i) determining the total molar flow rate of all gas phase CO withdrawn from the reaction medium of step (b); wherein the molar flow rate sum of step (i) divided by the molar flow rate sum of step (h) is at least 0.25 and not more than 0.80. [0065] This hydroformylation process can have a temperature difference of at least 1°C and not more than 12°C, at least 1°C and not more than 8°C and at least 1°C and not more than 4°C. [0066] This hydroformylation process can produce aldehyde formed comprising predominately propionaldehyde and butyraldehyde. [0067] In this hydroformylation process, the reaction medium can be comprised within a bubble column reactor. [0068] In another embodiment of the invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein said reaction medium has a highest elevation, a lowest elevation, and a maximum diameter, wherein said maximum diameter is at least 1.5 meters; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; (c) withdrawing from said hydroformylation reactor at least a portion of said gas phase effluent comprising at least a portion of said one or more aldehydes and non-aldehyde components, wherein said gas phase effluent is withdrawn above an elevation that is 0.25 times said maximum diameter below said highest elevation of said multiphase reaction medium; (d) cooling at least a portion of said gas phase effluent; (e) compressing at least a portion of said cooled gas phase effluent; (f) returning at least a portion of said cooled and compressed gas phase effluent to said multiphase reaction medium. (g) determining a time- averaged and area-averaged temperature of said multiphase reaction medium of said step (b) at an elevation that is 0.25 times said maximum diameter above said lowest elevation of said multiphase reaction medium; (h) determining a time-averaged and volume-averaged temperature of said gas phase effluent of said step (c); wherein the difference between said time-averaged and area- averaged temperature determined in said step (g) and said time-averaged and volume-averaged temperature determined in said step (h) is at least 1°C and not more than 16°C. least 0.25 and not more than 0.80. [0069] This hydroformylation process can have a temperature difference of at least 1°C and not more than 12°C, at least 1°C and not more than 8°C and at least 1°C and not more than 4°C. [0070] This hydroformylation process can produce aldehyde formed comprising predominately propionaldehyde and butyraldehyde. [0071] In this hydroformylation process, the reaction medium can be comprised within a bubble column reactor. Oxo Reaction L/D Ratio [0072] In another embodiment of the invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein the maximum inside diameter of said reaction medium is greater than at least 1.5 meters; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; wherein the time-averaged superficial velocity of said gas phase at the half-height within said multiphase reaction medium is at least 0.07 and not more than 0.8 meters per second (m/s); wherein the ratio of the maximum height of said reaction medium to the maximum diameter of said reaction medium is at least 0.8 and not more than 8.0. [0073] The time-averaged superficial velocity of said gas phase at the half- height within said multiphase reaction medium can also be at least 0.13 and not more than 0.8 meters per second (m/s). [0074] In this process, the ratio of the maximum height of said reaction medium to the maximum diameter of said reaction medium can be at least 1.2 and not more than 6.0, at least 1.4 and not more than 4.0, or at least 1.6 and not more than 3.0. [0075] In another embodiment of this process, the hydroformylation process comprises: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein the maximum inside diameter of said reaction medium is greater than at least 1.5 meters; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation, wherein the ratio of the maximum height of said reaction medium to the maximum diameter of said reaction medium is at least 0.8 and not more than 8.0; (c) withdrawing from said hydroformylation reactor at least a portion of gas phase effluent comprising at least a portion of said one or more aldehydes and non-aldehyde components; (d) cooling at least a portion of said gas phase effluent, wherein the enthalpy exchanged out of said gas phase effluent is at least 15 percent of the enthalpy of reaction provided in said reacting step (b); (e) optionally, compressing at least a portion of said cooled gas phase effluent; (f) optionally, returning at least a portion of said cooled and compressed gas phase effluent to said multiphase reaction medium. [0076] For the cooling step in (d), the enthalpy exchanged out of said gas phase effluent can be at least 35 percent of the enthalpy of reaction provided in said reacting step (b). [0077] In another embodiment of the invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein the maximum inside diameter of said reaction medium is greater than at least 1.5 meters; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation, wherein the ratio of the maximum height of said reaction medium to the maximum diameter of said reaction medium is at least 0.8 and not more than 8.0; (c) withdrawing from said hydroformylation reactor at least a portion of gas phase effluent comprising at least a portion of said one or more aldehydes and non-aldehyde components, wherein the ratio of molar flow rate of said non-aldehyde components in step (c) to the molar formation rate of said one or more aldehydes in step (b) is at least 3:1 and not more than 36:1; (d) cooling at least a portion of said gas phase effluent; (e) optionally, compressing at least a portion of said cooled gas phase effluent; and (f) optionally, returning at least a portion of said cooled and compressed gas phase effluent to said multiphase reaction medium. [0078] In these processes, the ratio of the maximum height of said reaction medium to the maximum diameter of said reaction medium can be at least 1.2 and not more than 6.0, at least 1.4 and not more than 4.0, or at least 1.6 and not more than 3.0. Oxo Reaction With Optimized Reaction Height [0079] In another embodiment of this invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst, wherein the maximum diameter of said reaction medium is greater than at least 3.5 meters and less than 12 meters, and wherein the total height of said reaction medium is at least 5 meters and not more than 26 meters; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation. [0080] In embodiment of this invention, the total height of the reaction medium can be at least 6 meters and not more than 23 meters, at least 8 meters and not more than 20 meters, or at least 9 meters and not more than 17 meters. [0081] In another embodiment of the invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; and (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; wherein said reactor of step (a) comprises at least one substantially cylindrical section and wherein the total height of all said substantially cylindrical sections is at least 6 and less than 30 meters in height. [0082] In these embodiments, the reactor of step (a) can comprise at least one substantially cylindrical section and wherein the total height of all said substantially cylindrical sections is at least 6 and less than 30 meters in height, at least 8 and less than 26 meters in height, at least 11 and less than 23 meters in height, or at least 13 and less than 20 meters in height. Oxo Purge Gas Reactor Process [0083] In another embodiment of the invention, a method of producing one or more aldehydes is provided , said method comprises: (a) carrying out at least one hydroformylation reaction comprising one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst in a main hydroformylation reactor to thereby form a first quantity of one or more aldehydes; (b) withdrawing from said main reactor at least a portion of gaseous effluent comprising one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and one or more aldehydes; (c) cooling at least a portion of said main reactor gaseous effluent to thereby provide at least a portion of condensed crude liquid aldehyde and at least a portion of cooled effluent gas from said main reactor; (d) compressing at least a portion of said cooled effluent gas from main reactor to form at least a portion of main reactor recycle gas; (e) introducing at least a portion of said main reactor recycle gas back into said main reactor; (f) separating at least a portion of main reactor purge gas from said cooled effluent gas of step (c) and/or from said main reactor recycle gas of step (d) to form at least a portion a purge gas reactor gaseous feed, wherein the concentration of olefin is at least 1 mole% and less than 20 mole%; (g) introducing at least of portion of said main reactor purge gas into at least one purge gas hydroformylation reactor comprising a liquid phase and a gas phase combining to form a purge gas hydroformylation reaction medium comprising one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (h) forming in said purge gas hydroformylation reaction medium of step (g) a second quantity of said one or more aldehydes, wherein at least 80 mole percent of all olefins fed into said purge gas reactor are converted to one or more aldehydes; (i) optionally, introducing into said purge gas hydroformylation reaction medium of step (h) at least a portion of additional feed that has not been previously introduced to a hydroformylation reaction medium and selected from the following group: olefin, molecular hydrogen (H2), carbon monoxide (CO), and inert stripping gases; (j) optionally, withdrawing at least a portion of liquid phase from said purge gas hydroformylation reactor of step (g) wherein the mass flow rate of said withdrawn liquid phase is less than 50 percent of the mass rate of formation of said one or more aldehydes in step (h); (k) withdrawing from said purge gas reactor at least a portion of purge gas reactor gaseous effluent, wherein the molar flow of aldehydes contained within said purge gas reactor gaseous effluent is at least 1.0 times the amount of aldehyde formed in step (h), wherein less than 50 mole% of said purge gas reactor gaseous effluent is subsequently fed to a hydroformylation reaction medium. [0084] In other embodiments of this invention, the concentration of olefin in said purge gas reactor gaseous feed of step (f) can range from at least 2 mole% and less than 15 mole%, at least 3 mole% and less than 12 mole%, or at least 4 mole% and less than 10 mole%. [0085] In this hydroformylation process, at least 88 mole percent of all olefins fed into said purge gas reactor can be converted to one or more aldehydes in said step (h). Other conversion rates can be at least 92 mole percent or at least 96 mole percent of all olefins fed into said purge gas reactor are converted to one or more aldehydes in said step (h). [0086] In this process, the moles of aldehyde formed in said step (h) can be at least 0.05 percent and less than 18 percent, at least 0.1 percent and less than 6 percent, at least 0.2 percent and less than 4 percent, or at least 0.3 percent and less than 2 percent of the moles of aldehyde formed in said step (a). [0087] In this hydroformylation process, the moles of molecular hydrogen (H2) and the moles of carbon monoxide (CO) in said purge gas reactor gaseous feed of step (f) can be, respectively, at least 2.0 and 1.5 times the moles of olefin in said purge gas reactor gaseous feed. [0088] In other embodiments, the concentration of olefin in said purge gas reactor gaseous feed of step (f) is at least 3 mole% and less than 12 mole% and wherein the purge gas reactor gaseous effluent of step (k) is subsequently cooled to produce a portion of condensed crude liquid aldehyde that is at least 1.03 times, at least 1.06 times, or at least 1.09 times, the mass of aldehyde formed in step (h). [0089] In another embodiment, the concentration of olefin in said purge gas reactor gaseous feed of step (f) can be at least 3 mole% and less than 12 mole% and wherein at least 92 mole percent of all olefins fed into said purge gas reactor are converted to one or more aldehydes in said step (h). [0090] In another embodiment, a hydroformylation process is provided wherein the concentration of olefin in said purge gas reactor gaseous feed of step (f) is at least 2 mole% and less than 12 mole% and wherein at least 92 mole percent of all olefins fed into said purge gas reactor are converted to one or more aldehydes in said step (h) and wherein the purge gas reactor gaseous effluent of step (k) is subsequently cooled to produce a portion of condensed crude liquid aldehyde that is at least 1.03 times the mass of aldehyde formed in step (h). [0091] In yet another embodiment, a hydroformylation process is provided wherein the purge gas reactor gaseous feed of step (f) comprises at least a portion of main reactor recycle gas of step (d), wherein the concentration of olefin in said purge gas reactor gaseous feed of step (f) is at least 2 mole% and less than 12 mole%, and wherein at least 92 mole percent of all olefins fed into said purge gas reactor are converted to one or more aldehydes in said step (h), and wherein the purge gas reactor gaseous effluent of step (k) is subsequently cooled to produce a portion of condensed crude liquid aldehyde that is at least 1.03 times the mass of aldehyde formed in step (h). Oxo Purge Reactor Gas-Phase Staging [0092] A method of treating an olefin-containing purge gas stream is provided, said method comprises: (a) feeding an olefin-containing purge gas stream comprising one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and at least one inert gas component into a multiphase purge gas reaction medium comprising a gas phase and a catalyst-containing liquid phase in a purge gas reactor, wherein the concentration of olefin in said purge gas stream is at least 1 mole % and less than 12 mole%; (b) forming in said multiphase purge gas reaction medium one or more aldehydes via hydroformylation, wherein at least 80 mole % of the olefin in said purge gas stream is converted to aldehyde before exiting said multiphase purge gas reaction medium; (c) withdrawing from said purge gas reactor at least a portion of gaseous effluent comprising at least a portion of aldehyde vapor and at least a portion of said inert gas component of step (a), wherein the residence time distribution function of said withdrawn inert gas component within said purge gas reactor provides a CMF value first exceeding 0.90 when the value of (time/average residence time) is at least 1.04 and is less than 1.95. [0093] The residence time distribution function of said withdrawn inert gas component within said purge gas reactor can provide a CMF value first exceeding 0.90 when the value of (time/average residence time) is at least 1.06 and is less than 1.68, at least 1.08 and is less than 1.55, or at least 1.10 and is less than 1.48. [0094] In other embodiments, a hydroformylation process is provided wherein said aldehyde formed in said purge gas reactor comprises predominately propionaldehyde and butyraldehyde and additionally comprising withdrawing at least a portion of catalyst-containing liquid phase from said purge gas reactor, wherein the mass flow rate of said withdrawn liquid phase is less than 50 percent, less than 45 percent, less than 40 percent, less than 40 percent, less than 35 percent, less than 30 percent, less than 25 percent, less than 20 percent, less than 15 percent, less than 10 percent, or less than 5 percent of the mass rate of formation of said one or more aldehydes in step (b), [0095] In these hydroformylation processes, said aldehyde formed in said purge gas reactor comprises predominately propionaldehyde and butyraldehyde and wherein the formation of said one or more aldehydes in step (b) can convert at least 92 mole% of all olefin fed to said purge gas reaction medium. In addition, the mass of aldehyde formed per hour in step (b) divided by the mass of liquid phase reaction medium in step (a) can be at least 0.1 and is less than 1.4. [0096] In another embodiment of the invention, a method of treating an olefin-containing purge gas stream is provided. The method comprises feeding a purge gas stream comprising ethylene and/or propylene with a catalyst- containing liquid phase to form a multiphase reaction medium disposed within a purge gas reactor pressure vessel having an upright orientation with L:D ratio of at least 8:1 and less than 180:1 under conditions sufficient to form, via hydroformylation, one or more aldehydes in said multiphase reaction medium with a per pass olefin conversion rate of at least 80 percent. Other ranges for the L:D ratio include at least 16:1 and less than 150:1, at least 24:1 and less than 120:1, and at least 32:1 and less than 90:1. [0097] A method of treating an olefin-containing purge gas stream is also provided wherein said purge gas reactor pressure vessel has an L:D ratio of at least 8:1 and less than 90:1 and wherein said pressure vessel contains at least one horizontal baffling means disposed within said multiphase reaction medium. [0098] In another embodiment of this invention, a method of treating an olefin-containing purge gas stream is provided wherein said purge gas reactor pressure vessel has an L:D ratio of at least 8:1 and less than 90:1 and wherein the average gas holdup within said multiphase reaction medium is at least 12 and less than 50 volume percent or at least 20 and less than 40 volume percent. [0099] In another embodiment of the invention, a method of treating an olefin-containing purge gas stream is provided wherein said purge gas reactor pressure vessel has an L:D ratio of at least 8:1 and less than 90:1 and wherein at elevations within said multiphase reaction medium having a maximum diameter D the minimum superficial gas velocity is at least 0.06 m/s and the maximum superficial gas velocity is less than 0.9 m/ or wherein at elevations within said multiphase reaction medium having a maximum diameter D the minimum superficial gas velocity is at least 0.09 m/s and the maximum superficial gas velocity is less than 0.6 m/s. Optimized Oxo Main and Purge Reactors [00100] In another embodiment of the invention, a hydroformylation process is provided. The process comprises: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby produce one or more aldehydes via hydroformylation; (c) withdrawing from said hydroformylation reactor a gas phase effluent comprising at least a portion of said one or more aldehydes and non-aldehyde components, wherein the molar ratio of said non-aldehyde components to said one or more aldehydes in said gas phase effluent is at least 6:1 (8:1, 10:1, 12:1, 14:1, 16:1, 20:1 and/or not more than 40:1, 35:1, 30:1, 25:1); (d) reacting at least a portion of said gas-phase effluent in a purge gas reactor under conditions sufficient to form a second quantity of aldehydes via hydroformylation; and (e) removing at least a portion of said second quantity of aldehydes from said purge gas reactor via an aldehyde effluent stream. [00101] In another embodiment of the invention, a hydroformylation process is provided comprising (a) carrying out at least one hydroformylation reaction in a main reactor to thereby form a first quantity of one or more aldehydes; (b) withdrawing an olefin-containing gaseous effluent from said main reactor; (c) contacting at least a portion of said olefin-containing gaseous effluent with a catalyst-containing liquid phase medium in a purge gas reactor under conditions sufficient to form, via hydroformylation, a second quantity of one or more aldehydes in said liquid phase medium; and (d) stripping at least a portion of said aldehydes out of said liquid phase medium produced in step c) using a gas phase stripping medium to thereby produce a gaseous purge reactor effluent comprising at least a portion of said one or more aldehydes stripped out of said liquid phase medium, wherein less than 25, less than 10, less than 1, or less than (10, 5, 1, 0) mole percent of said gaseous purge reactor effluent is recirculated back to said scrubbing reactor. [00102] In other embodiments of the invention, one or more aldehydes are stripped out of said liquid phase medium in said stripping of step (d) at an average molar stripping rate that is at least 80, at least 90, at least 95, at least 98, or at least 100 percent of the average molar rate of formation of said one or more aldehydes in step (c). [00103] The purge gas can comprise one or more olefins, carbon monoxide (CO), and molecular hydrogen (H2) in respective amounts such that the molar ratio of combined CO plus H2 to said one of more olefins is at least 3:1, at least 4:1, at least 5:1, or at least 6:1. [00104] In another embodiment of this invention, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a main hydroformylation reactor, wherein said multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (b) reacting at least a portion of said one or more olefins, said H2, and said CO in said multiphase reaction medium to thereby produce a first quantity of one or more aldehydes via hydroformylation; (c) agitating said multiphase reaction medium in said main hydroformylation reactor during said reacting of step (b), wherein less than 50 percent of said agitating of step (c) is attributable to mechanical agitation; (d) withdrawing a gaseous effluent from said main hydroformylation reactor, wherein said gaseous effluent comprises one or more olefins, molecular hydrogen (H2), and carbon monoxide (CO); (e) forming another multiphase reaction medium comprising a liquid phase and a gas phase in a purge gas reactor, wherein said multiphase reaction medium comprises at least a portion of said gaseous effluent; (f) reacting at least a portion of said gaseous effluent in said purge gas reactor under conditions sufficient to form a second quantity of aldehydes via hydroformylation; and (g) agitating said another multiphase reaction medium in said purge gas reactor during said reacting of step (f), wherein less than 50 percent of said agitating of step (g) is attributable to mechanical agitation. Retrofit Hydroformylation Processes [00105] Any of the previous processes can be conducted utilizing equipment that has been used in other hydroformylation, chemical, or industrial processes. Any previously utilized reactor in a hydroformylation, chemical, or industrial process can be used in any of these processes. [00106] In one embodiment, a hydroformylation process is provided comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein the multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (b) reacting at least a portion of the one or more olefins, the H2, and the CO in the multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; (c) withdrawing from the hydroformylation reactor at least a portion of a gas phase effluent comprising at least a portion of the one or more aldehydes and non-aldehyde components; (d) cooling at least a portion of the gas phase effluent to form at least a portion of cooled gas phase effluent and at least a portion of condensed crude aldehyde liquid phase comprising at least a portion of at least one aldehyde and at least a portion of at least one olefin dissolved therein; (e) separating and )compressing at least a portion of the cooled gas phase effluent to form a gas phase recycle gas; (f) returning at least a portion of the recycle gas of step (e) to the multiphase reaction medium of step (b); (g) returning at least a portion of condensed crude aldehyde liquid to the hydroformylation reactor of step (a); (h) contacting in a syngas scrubber at least a portion of a gas phase raw syngas feed comprising H2, CO, and at least one syngas impurity with at least a portion of the condensed crude aldehyde liquid of step (d) to form a liquid phase of improved crude aldehyde and a gas phase of improved syngas feed wherein at least a portion of at least one of the dissolved olefins of the condensed crude aldehyde liquid is transferred into the improved syngas feed; (i) supplying at least a portion of the improved syngas feed from step (g) to the multiphase reaction medium of step (b) or to the recycle gas of step (e); (j) supplying at least a portion of at least one olefin feed supply to at least a portion of the reaction medium of step (b) or to at least a portion of the recycle gas of step (e) or to syngas-recycle gas mix or to improved syngas; (k) optionally, separating at least a portion of cooled gas phase effluent of step (d) or at least a portion of the recycle gas phase effluent of step (e) to form a purge gas; wherein the process is completed in previously used equipment. Aldehydes Produced from Inventive Hydroformylation Processes [00107] In another embodiment of the invention, an aldehyde product is produced by the process comprising: (a) carrying out at least one hydroformylation reaction comprising one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst in a main hydroformylation reactor to thereby form a first quantity of one or more aldehydes; (b) withdrawing from said main reactor at least a portion of gaseous effluent comprising one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and one or more aldehydes; (c) cooling at least a portion of said main reactor gaseous effluent to thereby provide at least a portion of condensed crude liquid aldehyde and at least a portion of cooled effluent gas from said main reactor; (d) compressing at least a portion of said cooled effluent gas from main reactor to form at least a portion of main reactor recycle gas; (e) introducing at least a portion of said main reactor recycle gas back into said main reactor; (f) separating at least a portion of main reactor purge gas from said cooled effluent gas of step (c) and/or from said main reactor recycle gas of step (d) to form at least a portion a purge gas reactor gaseous feed, wherein the concentration of olefin is at least 1 mole% and less than 20 mole%; (g) introducing at least of portion of said main reactor purge gas into at least one purge gas hydroformylation reactor comprising a liquid phase and a gas phase combining to form a purge gas hydroformylation reaction medium comprising one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (h) forming in said purge gas hydroformylation reaction medium of step (g) a second quantity of said one or more aldehydes, wherein at least 80 mole percent of all olefins fed into said purge gas reactor are converted to one or more aldehydes; (i) optionally, introducing into said purge gas hydroformylation reaction medium of step (h) at least a portion of additional feed that has not been previously introduced to a hydroformylation reaction medium and selected from the following group: olefin, molecular hydrogen (H2), carbon monoxide (CO), and inert stripping gases; (j) optionally, withdrawing at least a portion of liquid phase from said purge gas hydroformylation reactor of step (g) wherein the mass flow rate of said withdrawn liquid phase is less than 50 percent of the mass rate of formation of said one or more aldehydes in step (h); (k) withdrawing from said purge gas reactor at least a portion of purge gas reactor gaseous effluent, wherein the molar flow of aldehydes contained within said purge gas reactor gaseous effluent is at least 1.0 times the amount of aldehyde formed in step (h), wherein less than 50 mole% of said purge gas reactor gaseous effluent is subsequently fed to a hydroformylation reaction medium. [00108] The table below shows various embodiments of the inventions. Any of the ranges can be used in combination with each other.

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_{r r} > g ^, _r n ^e _t o ^t d o o _, n a e ^t _c ^t _c e ⁱ _s a _, _l d a a ^d _y n w a m ^o _t e _r a e T ^> _i ^m e _te e y e e h d ^, _r x ^r _t ^l _t s h ^{r r} e no w O o a ^s u s e n n _l ^d o ^{t N t} _c ^m o e v _l ^{d i i} ^a _/ e e ^{a c} d g _i ^s _t a _i ^d e < u _sa a n w dy dy _i ^d n u ^{a n}i_{l ,} o a ^r n m _i ^{q l} g _i ^m i^o o^l _f a he he q l^{i e} _r o ^h _t o a e a _{r i} _t ⁿ ^t _ca _{s i} ^m _l ^d _l ^d e u ^c _ro ⁱ _t ^e _r d u _i ^{m o}c ^m e ^r a x a g a _f a_f du ^s _s &_t ^t _{c l} ^ay ^t _{s r} c _i ⁿ ae ^>, ^y _{r t} ^e n a _s d ^{m r} _r h g ^{o o r} o _i ^{n m} < ^e _{c r} e ^m _r a ^o _t _r ^e _{r f} p _i ^{e r o} _f g ^e _s ^> _i ^a ca m ^b p _, s p m e ⁱ _{r l} ^{e u} _t ^u ^a _t ^o _r b o dn ^o _{r t}n n ^, _ro ^v d ae r ^a _r ^e _r ^t _c m a² d e ed ^t _c e ^ru s ^t _s o e ^C _° e ^C _° u a ^yt g yh ^v n a vo mS ⁱ _{r t} ^m _/ p _,a p _, ss e ^r _ic nⁱ _y n ⁱ o e ^o _r _s p n _l ^e m m a e m e o e ^e n g o ^t _i ^m e ^t _i ^m _r e _l ^o _ti n n ^{c r} e x ^a o n ⁱ n ^{i )} _r me ^f h ^r d ⁾ _rhn c _i a _" m ^md m x , g a g _i ^{n p} g e w e _i ^u o ^e _{t i} ^{n m} _i ^{n m} _i ⁿ _t ^{v e} n e a w ⁱ o a _f o ⁱ _t _l ^tk ^% a e d ^/ ^r _tn ^d _y e _s e e m _l ^{e d} _y v ^/ o _s _l ^e ^R- n d b a s ^< s s a b ^{P l} _f ^e c _l ^e e h _r o h m o^e _r o e u z m ^r ne ^, _s ne ^>, _s ne Pk _la _{k, l i} ^m iⁿ o_f a b o m c e l^d o ^f m^/ e e r_l ^{d r} m^/ _rs _r ^{m e} n d a d a d _l ^o o n c ^t d o c_i ^{o i} n g n g n _,a ⁱ _{t i} n ^e _r ^a i^{c m o} n d _, n a o a_f ed h^/ a_f ed h^/ ^t n _t o a o c _t o e ^c _t o c _i _i ^{m e} _r us ^f _r > _l _{, i} ^{o e}c _i ^c ^{m n} % ^o yh _s _l ^{e o} yh _s _l ^e ^s a _c e a c ^m _{r i} ^{t l} _{t i} ^t e^l _{t i} ^t _i ^{n e} _f s e e p ^s _/ ^t _c ^y _{c i} ⁿ i_{f l} ^eo i^o _t e o i^o _t e o D ^r _" ^e _r e S ^o _f x E u x o E u x o E m _i ^f d _r p ^u _s m a^r _f ^e _r m _l ^e o m ^a _{r l} ^d a ^m ₍ ^a _{r l} ^d a ^m ₍ O C_, ^N _, , O s _s 2 _, ^, _s a e H _l ^s ed e s y dy n nⁱ _f n _, o 2 h h h ed % _l ^e a l^k O o l^c e l^d e l^d n - 8 % 4 _{- -} o a C a a a o c de d _r _i o i^v a d _i ^m e o d _{) i} ^d _{) s} _i ^d d r o_t x ^{t u} d u d u e a _r e c k ^c v _i ^q e l v _i ^{q l e} _f _r a ^m ee o o _i o o e ^t ^{r <} m ⁿ o ^, w n e _l m _i ⁿ e _l ^t _c a t _r o _r d _t e w ^r c _i ^a _te ^r i ^a _te e e a ^{p a} _r ^d d a d u ^{m d}u ^{m r} iⁿ e ^f e ^r e ^f i^d _{l i} ^m h x _i ^t _i ^{q l t} _{s i} ^q l ^{l t} _{s t} _l n e a w n ^{i y} n ^{i y e} u h ^{t a} _te _{r l} ^a _ta _l ^a _{t s} a i^q _l n ^m ou ^a _t ^e r ^{i m} e c_f ^a _te c _r p o ^t _s ^<, _r ^qi_l a m ^o m _f o^s _t u ^e _ta _l ^y o^t _{r i} ^m _ts n _ts n n q_l ^{i r} r n ^a _{t c} e xa a a ^h _{t l} ^y o aⁱ _ta _l ^y o aⁱ _{t r} e a o^t no n e oⁱ ^h _t _t a ^c e ^r o ^m _t ^r _{t t} ^r _t a _c p _i o ^h _i m_f <_, ac n a ac n a ^o e m ^t _c m ^m _{r s} ⁿ e _y ^o d _f e _i ^m _f e _i ^m e ^o c x ^r o a f o o^{f r} _ff d n _i ^{n o} cn a o c e a oⁱ ^o _t n o n o ^t _r ^r e e a ^m _r oⁱ ^o _t ^c ₍ ^m o^{i c} ₍ ^m d e e i _l _t ^d _y ^oi_t ^r ^{a r} o e^f a ^> _t r _y ^, ₎ a ^< e r _y ^, ₎ ^{f h} _t h o ^t ^a _r h _t e s e l ^a ^d _r m s_i ^{u a} _r d _t _s y n ^b d _t h e d c e _i ^e _l n ^b d ^s _{t l} e d _i ^e n a ^{n i} n d p c e _l a^{t o} e sa ^a s d s e n _i y a o ^d p n _i ^d pp _c i_t ^m ^h a _l ^d o _i ^{v u} o _i ^{v u} a p _r m b m _r m a ^c ₍ d _s ^c ₍ d _s ^e _r O ^o _f o ^t _cae R₎ _i ^{n ,} a ¹ _s e ^t _{c s} ^M _(t g u e n d s a 1 R ^o _i _r ⁿ p m a - % 4 _{- - -} ⁴ _. 0 ⁰ _. 5 ⁰ _. ⁰ _. 6 82 ⁵ _. ⁰ _. 4 61 ssa _,d _, - - - - % d ^{g g g g} d ^m e e e e e n a a ^e _f ^e _f e d e d e d a_{l i} ^m mⁱ _{x r}o _r d o y y d h h y y d y h h y h e x s a a ^% a ^t _c ^t _c e e l^d e e h l^d e e l^d s ^m e ^m _i ^m ae ae _l ^d ^{a a} _{( l} ^d ^{a a} _{( l} ^d ^{a a} ₍ v^{< <,} _s xa ^r d ^r d _{(, , (} a _{, , (} a _{, ,}a n ^, i^o _s t d a n n a g ^{m 2 2} _i _c e _t <_, n ^m _i ^m a i_i ^m a i_i ^m ⁱ n _i ⁿ x ^m a _i ⁿ x ^m a _i ⁿ xa a ^e e _f e^l _t 2 e ⁱ e r _s u 3 n n ^{m m m m m m} ^y _r ^h _t a ^> o ^g o_{r s} a _l ^ey _l ^ey ^{, <, >, <, >, <,} e o g ^h _t ^h _{t s}s _ss _ss _ss _ss _{) s}s ₎a ^t _{c t} a a a a a ^L a ^Lm ^b s ^g _r a n e_, e_, m mm e u u sⁱ _r ^p e ^r e v en e ^m n _i ^d ₎ ^m g _i ^d ₎ ^m ₎ g _i ^{d L m} ₎ e _i ^{d L m} _i ^{u m} _i ^u e _i ^d d _i ^d dS _s p n ⁱ n n ⁱ n n ⁱ _l ^e _l ^e u n y y ^qi_l ^k u qi_l ^k u qi_l s u a ^qi_l s u e u e a ^qi_l ^qie mg o o o o p p _r es _r es _r h _r h _r ^m _ln c m ⁱ _t d _r ^md c ⁱ _t ⁱ _t ^o _r ^o _r e a e a e ^p e ^p e ea _" ^R- e _i ^u c c p p ^p h ^p h ^p n d ^a _r ^a _r ^a _{r , ,} ^{e p e p e} _i ^{d p} u ^e _i ^{d p e} _t ^{p e} _t u ^{e a} _r ^{e a} _r ^e _r o e ^{f f f} n s_s ^s _s ^s _s ⁱ _f nⁱ _{f t}a ^r _i ^d _t u a ^r _i ^d _{t t t} e _t e u a _i ^q _l a _i ^q _l a a a au z s _r ^m a a a _l ^e ^o a _l ^e ^o a n ^q _l ^r ⁱ n ^q _l ⁱ n ⁽ _/ ^r ⁾ _r n ⁽ _/ ^r ⁾ _r n ⁽ _/ ^r ⁾ _r n ⁽ _/ ⁾ _r _l ^o o n c ^t _{c i} ^{o m m m} % _i ^mx % _i ^m oⁱ _t ⁽ _/ ⁾ oⁱ _t ⁽ _/ ⁾ oⁱ _t ^h _/ oⁱ _t ^h _/ oⁱ _t ^h _/ oⁱ _t ^h _/ _i ^s a ^t _c e a _i ^du _i ^d i^q u _i ^d i^q u _l ^e a _l ^e _i ⁿ ca _r _i ^h ^q o ^m o m e _/ ca _r h^g e _/ ca _l ^e ca _l ^e ca _l ^e ca _l ^e ^g e o e o e o e o D ^r _" ^e _{r l} % _{l l} m < m > R _k R _k R m R m R m R m _i ^M ₍ e _t gn % 8 5 8 s a 1 R _{- - -} ⁵ _. % % % 0 07 % 2 06 % 4 07 ¹ _. 0 ⁸ _. 0 ⁵ _. 0 ⁹ _. % 0 0 % 2 07 - ,_r ^, _r o^t o c ^t _c d _r a a n o a e a_l ^t _c a ^> a _i ^m _i ^m a a a a ^, _r <_r e ^{r r} es e _i ⁿ xa s ^r _i ^m _i ^m _i ^m o^t _c o^t _{c l} ^% _l ^%o e d ^{m m} _i ⁿ xa _i ^m iⁿ xa ae a o e ^{v v} ^v n n ^{2 >} i^o n ^, _r ^<, ^t _c ⁱ g o ^o _r ^{m m m m s} _/ ^s _/ ^r i ^r 0 0 2² t ^o _t ^{>, <, > <} m_, m_, ⁿh _i ⁿh n na a ⁱ _t o ca c _{r r} ^, _r ^, _r a a _i ^t _i ^t a h % a o^t o^{t o} _t ^o _t e _i ⁿ _i ^m _i ^{w w} h ^{t t} _,a r p i ^a ^p _r ⁱ _t _, ^a _r ^e _rf ^e _{r c} _f a _c e a c c ^mx g g ^e ^e a a _i ⁿ a ^{U U e} _{r r} o o _i ^m ^y _r ^h _t a o ^r _t a _, s _i ^m a o_t o _{rf rf} ^e _r ^e _r ^m u _tu o o _fo _fo > ^m ^< mu mu ^{m m} _i ⁿ n ^mm ^b d s _i ⁿ _i ^{m o o} _t u _t u _{t t ,}g g m e es u s a ^m _i ⁿ e e > ^m s ^{o o} u u ^{U U} _i n m^{i i} > x _i ^mx g g _,g a sa e e ^o e ^{o y} _t ^y _t a a ^{E E}S ⁱ _r ^e _{r ,} > h h s s s es _ic _ic ^{m m} n no ^E _s p c n e m g o _i ^{n i} _{f ,} n p p a i c m _{r l} ^e _{f s}a _s h a a p h a a s p h h _l ^o _l ^o o s ^{p p e} _v ^e _v ^t o ^t _i ^ot _c ⁱ _t n c oⁱ _t m m a ^a _r cna _{" i} ^o _f o_{: l} ^{e g g} a a s s _{s s} u u _f ^{r f a} _r ^R- e ^u n d ^e e _s O ^o _: a a 2 ^{% % g g g g} ag ag _i ^m _i ^m _s s a a % ^fs_r o e _{l l i} ⁿ _i ⁿ % ^g _, au z s _r ^m s n a ^C d ^H _l ^e _l ^{e % % % %} e de o o _l ^e _l ^e _l ^e _l ^e _i ^a _i ^{a g} _, m m m a m a ^g e_l ^o o ^g c ^t _{c i} ^o d ^r _r ^r _r ^{m m} o o o o _i ^cf _{r i} ^cf _{r f} o a _f a o _i ^m u _i ^m u _i ^m i_i ^m x g a_i ^s a ^t _c e e a d ^e _f ^e _f ⁿi_f ⁿi_f ^{m m m m} ep ep i^o _{t i} ^m iⁿ i^o _t x _i ^{m n} a _r a _i ⁿ m x e D ^r _" ^e _r d a ^e _r p ^e _r p _l ^e o _l ^e O O 2 2 o C C H H ^u _s ^u _s ^a a r m ^a _r m m > m ^m < v a ₎ _i ⁿa ¹ ^M ₍ e _t gn %s a _y % 1 R - n a 05 2 ² _. 0 ⁰ _. 4 ³ _. 0 ⁰ _. 6 _{- -} ¹ _. % 0 81 52 n oⁱ _t ^< a a ₎ a d _i ^t m _s a _s a _s a _s a _i ^m _i _i ^{n m}x ² _^ n g _i ^u ^a _l ^a _l de ^g P ^g ^d P ^g ^d P ^g a a ^D _* P n i^{m oi} _t ^m ^> m <⁴/_i e ^a _t ^{m d d} x c ^, _x ^, _x p ss o e ^t n ^{* * *} ^v % _i ^ot w w w a ^* w a a a c o ^{m e} _r ^{r r (} a n _, ^l _f a o^l _f a o^l _{f i} ^m o^l mⁱ _{x t} ^t _s ^t _s ^e _r _i ^o a a _{s s} mⁱ _{s i} ⁿ _f _s a ^<, _r s ^{1 1 A} ^t % c_, e a a _i ^mx ^r a _i ^m a _x a 3 ^g _i ^{n m} a ^m o ¹ ^t o ^o _t ^o _t o ^t e _i ^m a ^m _/ e ^g a e ^{g > g <} _ca ^t _x ^r _x ^r d m ^m m ^m e m e e s m es e ^r m d d e r x ^m w _l ^{u >} _l ^{u <} _l ^u a _l ^u a ^t _i ^u n n d ^{2 2 m} _r ^y _r ^h _t a a o ^{m <,} _r ^k, _r o m o m o h o h _s V_{( i} ^ud ^V _{( i} ^ud ^V ₍ ^{p V} ₍ ^{p 1} e o o o ^f ^{m i} _t ⁱ _t em ^b s <_, e e _r e _r e _{r i} ^d _{r i} ^{d o} _t n a ^r a ^r dym e u s g w i o wo e r ^{E p p} w ^m e ^m e u q e u q _ro _i ^{o e} _t ^e _t h o n w n w i_l w i_l ^t _c ^t _c a a eS _s p n n n ^p _i ^o o p _i ^o o p _ro o p _ro a a ^{r r} _l ^de m o g o ⁱ _t o c ⁱ _t oⁱ _t n ^t o _c ^t a no _c ^t a no _c ^t a no _c e e a n n a a ^{r r}n c m ^at_i ^at_i ⁱ _i oⁱ _t oⁱ _{t r}ha _{" i} ^{u a} _r ^f g g _ta e ^{r i} _ta e ^{r i} _ta e ^{r i} _t d d ^m a e ^r n n xa a a ^/ _s a^R- en d a a e s _i ^t 3 _i ^t 3 _i ^t 3 _i ^t 3 ^{2 2 m} a e _{l r} ^m _{r l} ^e _i ^m a _l g g g g oⁱ _t oⁱ _t ^m o o o _i ⁿ ^e _r o u z_r ^{m g} p a y_{t i} ^c _i ^{c a m} ^k _/ _r ^{a m} _/ ^{a m} _/ ^{a m} _/ a a < ^{f f} W ^k _r W ^k _r ^k _r ^{r r} _, e e ^m- ^ms o n eg _r n n a o ^k o ^k o W k o W e e m d d g >_l ^oc ^t _{c i} ^o o ah _r ah _i ^{m , , , k,} _i ^u y y ^k _,2_i ^s a ^t _c ^a e^r a _re ^t _c e v a c e e w c o e x ^w ₎ ^w ₎ ^w ₎ ^w ₎ m m a sa w s o w s w s w _l ^u _l ^u d h h o l a ^ol a ^ol a ^ol ^{o o} e e l^d e l ⁱ ^d _ta ^m D _{" r} a ^e _r m p m m g _f g _f g _f g _{f v v} m a a R _i ⁿ _i ^M ₍ _t gn s a 0 ^D 1 R 01 ⁵ _* _. 1 21 6 5 03 62 ⁰ _. 1 ⁵ _. 6 ² _. 1 ⁸ _. 0 ⁴ _. ⁶ _. 06 0 0 ⁶ _/ 0 1 8 0 0 3 1 9 1 H ) < 2 d _{^ ,} R n ^D _* C d 4 B n a_l /_{i r} a a ^{<, <, 2} e p ss ⁽ o^t a _c P Pk _{s s} n ⁱ a k_{, ,} a g a g s e v^e _r e ^r a a l^e _l ^{e e} _s n i ^A m m o^t o _{, , i} ^m _i ^m a mo _i ⁿ xa ^c _y ^c _y a g c ^t a a d _i ^m _i ^{m m} _t c o ^{m m} e c r e d r ed e e m m r^m _{, , i} ⁿ xa a a a > < m ^b _{, ,} ev ev d^y _r ^h _r _t o a a ^{m m} _i ^{m i} _{x i} ^m e x v P P o o ^a ga o ^f ^b e _i ^m _i ^mx ^{> <} _i ⁿ a a o d_r d_r b b _i ⁿ d _i ⁿ a ^, ₎ ^, ₎ ^{m m m} b a^o _t ^o _t ^a d ^a d dm s y m m R R ^> u u e e e e u s h m _, ^{m m < < D} d d ^e _fS ⁱ _r e _,a a m_, ^> m _s p _l ^d _i ^m a ^, _{) ,}a ^< C C ^, ^, ₎ a ^, ₎ ₎ ^{B B} a ^, m m _t ⁱ _i _x n ^b _i ^b o ^r _t ^r _t d a d _roe m a _i ^m x m m m _i ⁿ _{t i} ⁿ _i _t ^m iⁿ _i ^u _i ^mx _i ^{u e}- ⁱ _ta _i ^sd _i ^{s a} d O O ^fLg o _{r i} ⁿ a _i ^m c m h^/ a _i ⁿ _i ^{u i} d _xa _i ^u d a d d h h ^m e ^m e ^L c n o ^l _s a _s a C_r C_r T ana _" ^R- e _i ^u _s n d _l ^e _i ^{m m} > ^m < ^m e ^m e _i ^ge _i ^ge >_, ^m n <_, ^m oⁱ _t d ^{g g o} % l^{e o} % ^L _i ^m m _i ⁿ ^e e o xa _, _r o u z ^{m m}- ^m D _, _r D > r_, ^m < l⁾ n i^o _, ^{m h h} _l ⁾ o l⁾ n i ⁱ ^o eg eg e s _{t l} ⁾ n e o e d d 2 i_t ^a _c ^e _f e e H o 2 _l H ^eo ^o _rs _ro n g < _s c ^s _s c o ^l _s e ^f _l e ^f _{l f} m_{, f} m ^{f m} n ^> , ^e _t ^e _t e s^t _c e s^t _c ^al_l ^al_l e a e e a e _i ^t a a a a ^o n a ^o n _,a o^{i ,} _r _l ^oc ^t _{c i} ^{ot k} e e _s _i ^s a _c o 2 a _s a u ⁽ u ⁽ v ⁽ _r ⁽ v ⁽ _r ⁽ x e^g _i ^m no n _t o o _i ^o _i ^m _i ^o _i ^{m a t} e a ⁱ _t ^m m D ^r _" ^e _r a R _i ⁿ _i ^a m D _i ^{a e} _v ^e _r ^{( e} _v ^e _r ⁽ D ⁽ L H ⁽ L H ^H- ^D L ^H- L ^D _/ L _/ ^D H ^D _/ L _/ s _i ⁿa xa ⁱ _t ⁱ _t ^t _r xa ^t _{r i} ⁿ _{v c}a H a g m m p o p o o p m o p m _l ^e e ^e _r _i ^M _(t g ⁹ n _. ^D 0 _*0 s a _* 1 R H ⁴ _. 0 ¹ _. % 0 05 % 1 _{- - - - -} 0 0 6 41 61 ⁴ _. 0 -_r m ^e _t _i ^{s n} ^o > < d d _{, t} a m n _{) )} < >_, ^>, _r n n > e ⁱ e gv gv _, m o a_l ² _i ^m e _i ⁿ _i ⁿ _, D e ^d v e m v _l ^u _s ^a e _t ^/ _t ^a _t m /_t ^{C C} _{° i} ^u _i ^{u t} _c d d a s _s ^m n o ⁱ _r ^{( (} _{° ,} e ^e _r s e > b ob o e n n _, a e _t e s ^o a ^{v s} e e ^C _° ^C a ^{m m} s v a _,e d a a n i ^g c e _{r i} ^m _r g n h h _{, °} a _{, i} ^m _i ^m n n 1- o a s o^t x t d a o^t _i ⁿ i ⁱs ^{w w} a _i ⁿ xa _i ^o _i ^ot _r _c e a d ^f _r a _{c c} ^d u ^b R ⁹ _. ⁹ _{. i} ^m ^, e d _t a x _i ^m e ^m ae a _i ^v d T ^{0 0} a _i ^{n m} > ^{m t} _{c c}a o < ae e ^t _c r^{a s} e ^y _r ^h _t g m ^l _t ^r o ^< d ^r o _i ^m b S u ^{≥ ≥ m m} _{, ,} ^{r r} ae iⁿ _i ^u u ^t e ^t _i ^{n s C} F F m m d ^o e e ^{m .} _s d < e M M _, > _l ^e _l ^eg ^r d s _,s _i ^ud _i ^u g _i ⁿ na o d e s dy ^m _r d m ^b o y > ^v _i ^x C s e _, C_, d d e de _i ^{n s 2}m e ^e _f ^m a g he ^f h e e _,d D D D e e ^s ^{m m} n n ⁱ n ⁱu s _r ^e _l ^d _l ^d e T ^m l^y T T ^e _f _r ^e _f _r n n ^{i e e}S ⁱ _r ^o _{f v} w l _s p ob ^oe ^{a d} _y d h ^{a m} _r ^{R t} _c ^{R R} oⁱ oⁱ ec _{c c} e d o e e e e ^o _tc ^o _tc _tc _tc n n ne m L a a b ^e _l ^{e f} sa ^f _r sa sa a a a a ^e _r ^e _r ^e _rg o T _{r r s} ^d n c m ^L _i ^{m e} _t ^e _t n a _s e f n d h e h h ^e y p p p p _rf ^e _rf ^e _rf ^e _rf e^f e f_i ^ff_i e^ff_ia _"e _i ^u m x _i ^s _i ^s ed ^o ed h _{s ,}e _{s s} ^{o o o o d d d} ^R- n d a ^e _r o e ^o _r ^{f m} me m n n e a e ^o u z _c ^{% o} _{c l} ^{d g} m a a l ^e ^{u g g} _r ^e _r ^e _r ^e _r ^e _r ^C _° ^e _r ^e ^u _r _t ^{C u} _t ^Cs _r ^m n no ^<, _r ^d _l ^d _{l f} ^, _{r f} a_f de o de d ^u e a _t ^u ^a _t ^u _t ^u _t ^u _t _r ^a _r ^a _r ^{a a} _{, ° °} a ^a _{r ,} ^a _{r ,} _l ^o o c ^t _{c i} ^{ot i} _t o^t a n a a o ^e _t o r ⁿ _{r i x i} ^s _x ^{o r} _r ^v _l a ^r _r a ^r _{r r r} _i ^m e e e e e _i ^m e a p e a p_i ^s a _c e a ^a _{v c}a ^e _t ^{e m} iⁿ l^u _f m l^u _f ^{% e} _f u ^e _{f i} ^m i^{n e} _f x p p p p p a m m m m m x _i ^m _i ^m a m _i ⁿ m _i ⁿ D ^r _" ^e _{r l} ^e e ^e _{r i} ⁿ _t _i ⁿ m ^e _r e d ^e _r s a ^e _r p q e ^e _r p m ^e _r p m ^e _t ^e _t ^e _t ^e _t ^e _t m e ^t m e ^t m _ia ^M ₍ e s ^g _t gn _l ^o a c h _i ^d ws a _l 0 1 R _{- i} ^s c d x u ^o 0 e ^q _l ⁱ 03 8 1 e b 07 2 ⁵ _. 0 ,_{t l} ^o _t < h a ^{n i} h < d ^, n _{r i} ^{g n} _r d g o e ^e _t u _, m ^e _t > n a_l ^t _c ^h x , u ^o _{r ,} o m e _l ^o _r h L ⁱ _t e a < s e s ^r _i ^ud _{s f} a o w _f ^t ⁱ g a T ^a _re e ^t ^v _s e a 1 ^g p ^a _r p _i ⁿ w P ^L ^e ae n - m_f n ^e _t d ^e i ^s h _t ^k _v s _l ^o m o ^d ^o _r ^o _i ^ot _i ^t c n ^f b n w n e u ^{a i} ^t _c ^o _t a c n a oⁱ _t ^a o r ⁱ ^f _t e oⁱ _t s _i ^m w ^y _ti _l ^a _r vo ^a _r ah xa ^a _r c e ^r e ^r ca o e^y _r ^h _t d v a b ea ^{p m d} _l ^o e eva o m ^b n s ² _s ^r _{f ,} e ^a e _s a e _i ^du < _i ^d _s a _i ^{d d} i^{m d} _i ^q _{l ,}p ^{s d} _rm e n ⁱ ca _t u qi m _i ⁿ m _r ^e _{t i} ^d au s e d ^f _r h _ll ^o _r ^{f o} _r ^f o^{t s} u wS ⁱ _r c o us _i ^ge a ^{n m n} _c g _i ^{q l} n_s p n ^h _{t l} e m ^{h n} _{r r} >_{, r} a eⁿ _l ⁱ g wog o ^e _r e a e e ^u _t p ^u _{t r}n c m e^ff ^{m n} _r b ^t _x ^e _r o ^e _{r f} o ⁿia _{" i} ^u _i g ^e _t u ^t e _s ^{o l} _s ^o o _l d o _l p ^c o _i ^a ^R- en d ^{d n} _l ⁱ _i ⁿ _{l f} a g a e ^o _i ^d c _i ^cf ^e e _r o ^e _r ^C _° oo _r o _i ^ax ^o g n ^% _f ⁿi_l g_f g _r u _f a^{d q} _r i_l ^o e ^s _/u z s _r ^{m u} n _t a_{r ,}a c _f _{r t} ^a _l ^oi _t mu ^o o n o ^o _l n u ^e _{r l} n m a o _, pu ^m a_l ^o o c ^t _{c i} ^ot e _i ^m o n a t _l ^{a n a} _{r i} ^m oⁱ _t ^c oⁱ _l _t e us ⁿ _r ⁱ _t a ^s _i ^m _i ^s a _c p e a m xa _ca o _r e_t ea xa a _i ^d s s ^a c u a s ^{e e} _{t v i} ^m iⁿ _i ^du xa D ^r _" ^e _r ^e _t m ^e _r o C ⁿ _I % e d m _l ^o _i ^q _l c l^{o e} _{v r} P x e _l ^e e m _i ^q _l m ₎ _i ⁿ a ¹ ^M ₍ eg 0 _t n 5 0 s a 1 R ⁵ _. 2 ⁵ _. 0 D 3 ² _. 0 0 8 06 - ⁰ _, 0 6 _{- - - -} 61 ,_, a ⁿ _r ⁿ _r a a a i^m >_, ^u _t ^u _t a Pk ^P _k, ^P _k a , P ^e _r d x n a ⁿ _r e ^r e ^r Pk _, a a ^{k u} _s a _, a a s l ^{m u} _{t i} ^du _i ^d m m u a _i ^{m i} _x ⁱ _{x i} ^{m e} _r e _{r r} < e s e s ^t _f e^t _{f ,} ^r qⁱ qⁱ a _i ^m a P _i ⁿ a a _i ^{n p,} _t e v^{a a} L d ^{l l} ^s _t T n g g ^P _{k i} ⁿ k_{, ,}d ^m _,d ^{m m m} e^l _t _i ^s _ti ^{L a} _i ⁿ _{l i} ⁿ _{l ,}a ^m > a e >_, e <_, <_, >_, u n i^o u u ^e ^t _{v l} ^yp o o _i ^m _{, i} ^m ss d s n s dn _, ^t _{s ,} ^t _s o s_c dn dn o p oc oc _i ⁿ m xa ^e _r 2 ^e _r 2 d 1- d 1- a a o o b a u s _r o _r a o^{m e} _r ^m p _-t p _-t ^e _{s t} ^e _{s t} ^g e ^{r c c} _{i t}c _tc ^t _s mo e^l mo e s e e^l _t s e e^l _t su^y _r ^h _t w w ⁿ _r ^du a a ^>, _{t s} ^<, _t ^c _t u ^{c l} _t u _r p u _r p u _i ⁿa o ^a _r ^a _r ^u _t m ^{b d} _i ^q e^r- e^r- e a e T o T o m o m o m s e ^s _/ ^d e e ^s _/ ^{r l l} g w ^m w ^m _tuo ^{g l} _t e u O _s a O _s a o _sa o _sam e _i u s ^{d c} g ^c g m m_{, i} ^d m _i ^du ⁿi_{l l} ^oe ^, _{r l} ^oe ^, _{r s} ^g o r_s ^N ^s g ^N g o i _r ^s i _r S _r S _r ^t _tS ⁱ _r ^s a ^s _, _s p n ⁱ m n ⁱ a _i ^{q l} oo b_l o^t _c b_l ^o _tc a u a a g p g sa ^o _tc sa ^o _t ^I o c sa ^t _c ^Is o^t _c ob ae m g o ^y _ti ⁱ _x ^y _{ti i} ^m g ^c a i n c ^{n n}i_l n n a o e a r n e ^r _ro _{r r} t o o^{t g} ae ^g a a a a _r e ^g e ^{r g} e ^r o^t Pka _" m c a c i^u _l ^{o m} _l ^{o m} o e o e ⁱ _t o p w ⁱ _tp w _{c t} _l ^o a n _c d e a e ^r d ^r d e ^t _s e ^t _s e d d n e d _{c ,} n a a^R- en d e < e > c ^w ^e e ^v _t o _l ^o o ^e _r v ^e _r _r o _{l ,}n ^v _{l ,}n _f o e b _f e o^b _f e o^b _{f f f} ^f 1 e e _f ^f 1 e e _f ^f 2 e e _f ^f 2 e _f ^e _{rf i} ^mx a ^{o o o g} _r ^{o g} _r ^{o g} _r ^{o g} _r ^{o o} au z s _r ^m _i ^{a o} _l o n _i ^{c i} _{t i} ^a i^{c oi} _t no n m i oⁱ _, n a n a oⁱ _i ^m o _i ^{m e} _r ^e _r ^e _r u ^e _r u ^e _r u ^e _r u ^e _r ^e _r ^m ^oc ^t _{c i} ^{ot f} _r ^a _r ^f _r ^a _{r t t t} ⁱ _t x u u u ^p u ^p u ^p u ^p u u ^< _i ^s a _c e e e e a p a p a ^a _v ^a _{v i} ^m i^{n a} _{v i} ⁿ m ^a _v a ss s s n s n s n s n s s e s e s e e s e s e s e s s p e u u D ^r _" ^e _r S e d S e d _l ^e e _l ^e e m _l ^e e > _l ^e e ^m < _r p _r p _r h p w ^e _r h p w ^e _r h p w ^e _r h p w ^e _r p ^e _r p ^o _r d ₎ _i ⁿa ¹ ^M ₍ e _t gn s a 5 05 0 ⁴ _. ¹ _. 0 0 5 00 0 1 R 1 5 7 0 2 7 - 3 8 1 0 23 03 08 e_r ^e _ru ^e _{r i} ^c u p ^{< >} d ^u _s s o a % s ss _{, ,} ^r _t % ^, _r n _i ^m _l ^e _l ^e e ^, _r ^< b e ^, _r ^> e ^, _r e _s n s ^e _r ^e _r e _i ⁿ o b b b a a_l ^e _r p p de de _i ^s o m m b u b b b r ^u _r ^u _r ^u _r g es ^p s ^, _{, ,} _{t i} ⁿ _i ^{n m} _r ^m _{r l}a ^m ^> _{, ,} a a ^{c c c c n} _ys e e _s o o e ^, _{r s s s s} l _s ^{f f} d v _tu a _i g ag e ^m _i ^m _{s s s s x} d e ⁱ o o ^t _{c i} ⁿ xa ag ag ag ag ^e n o i^o _{s s} y dy ^t a n n n n e t _s u u h h e ^{m m} y y y y sa c a g _i ⁿ _i ⁿ e l^d e ^e _v _l ⁱ _t ^{rt > < s s s s} h ae ^r s m m u _tu _tu ^{a d} _l ^a _si ^{r e} _s ^e _s g i _l ^{a n} g e _i _l ⁿ g e^e _r ^f a a _i ^t _i ^t _i ⁿ g i^t _i ^{n p} i^t _i ^du^y _r ^h _{t i} ⁿ o o o o _yc ⁿ i h p h p x e x e x e x e _i ^qa o ^m _{s s} ^{m m} n P _i ^d _i ^{d l} ^b a a _{- -} d ^e _s ^e _s ^e _s ^e nm s m m e o g_r g_r gk gk _i ^e _/ u i^q u _s _i ⁱ ^q a a a a nu s ^t _t o a P o a _r a _{r i} ^cf _{f r} o ^{l l} h h h h ⁱ _{f r}S ⁱ _r ob ^s _s k ^s _s Pk e e a e ^t _c n n ^{p p p p} _l ^e _l ^a _s p _r a _{, ,} ^p _i ^{m p} _i ^m _r a _i ^o _i ^o _i ^d _i ^d _i ^d _i ^{d o} oe m o P ^e _r a ^e _r a ^k _{r i} ^{n k} _r x o ^e _i ^t _i ^t u u u u dg o ^t k p _i ^m p o o a s _r s s ^qi_l ^qi_l ^qi_l ^qi e ^mn c m _ca _, m x m _i ^{m m} m s e _s op op _{f f f lf l} ^v ma _"e _i ^{u e} _r a o a o _i ^{n w} n >_, ^w < _r a ^{o o o o o} _s p^R- n d c c n ^e e _{f i} ^m _f _r ^m _f ^m _, p ^g mo mo ^{e e e e} _i ^p o ^o _i ⁿ _i ^o _l ^e _i ^o _l ^e m e ^{c c} _r ^C _{° r} ^C _r ^C _{° r} ^{C s}d ^<u z s _r ^{m o < o > s} n ^e _r ^{m e} _r e ^e _r e _s o ^s ^m _s o o c^g _ru e e ^u _t ^u ^a _{, t} a _°, ^u _t ^u ^a _{, t} a _{°, f} ^, _re_l ^o o^t _i ^o us ^> u s s a u s ^e _{r -} ^e _r ^m- s a p g p g _l ^e d d _r a _r a _r a _r a o c ^p y y e _i ^m e e _i ^m e _s bc _c ^t ^s _c s p s e ^e s ^{e k} _/ ^k _/ n h h p p _i ^m p p _i ^m _l ^e b_i a a _rc _r m m ^y _c ⁱ e e m xa m _i ⁿ m xa m _i ⁿ o ^u _r D ^r _" ^e _r ^e _r p ^o _r d ^e _r p _i ^{n e} _r c o p _i ⁿ C ^W _k o C ^W _k ^e _r P d _l ^d a _l ^d a ^e _t m ^e _t m ^e _t m ^e _t m m ^c _s _i ^M _(t gn sn 8 00 s 0 0 0 5 0 5 5 0 a a 1 R ^r _{t - -} ⁶ _. 0 21 ⁰ _, 1 54 55 - ³ _. 0 ⁸ _. 0 ⁵ _. 0 ⁹ _. 0 - ² _. 0 ⁸ _. 0 , ^> _, ^< >_, d m ^, _r n n e ^o _t m ^, _r e ^o _t ^I ^a _l ^t _s ca ^t _s c ^o a _tt a _l e y _i ^a a P _l a a a a a a _i s ^s e y s d ^{rt s} e u _i ^{m t} _r ^k _i ^at Pk _{i i} ^m a _i ^{m m}x s d ^{rt O} _i ⁿ a _{, r}a _, ^m iⁿ x _i ^m _i ^mx _i ⁿ a e n n _s s ^p a a a v a ¹ a ¹ _l ^{e m} O _i ^p _i ^{m m} _i ⁿ a ^{m m} n _i _i ^g ^{o l o} _{t i} ^{g o} _t o > a C ^mx 2 H ^mx > < ^{m m} >_, <_, t_c e _r ^l o e _ro m _, ₎ oⁱ a _t ^P _{k l} ^l a a _ll a _{, ,} > a _i ⁿ _i ⁿ _, <_, ⁿ i ⁿ i m ^t m ^t 2 _, ^{m m} _i ⁿ _i ⁿ s s e a _ca a _ca ^H a ^r a m s^< m ^< O O C C 2 2 _l ^eo _l ^e ^{r s} e ^s e ^{+ i} _{/ i} ^m g ^, _tu ^{s ,} _t ^/ _t ^/ _t H^/ _t H^/ _t o^y _r ^h _t e ^e _r ^rd ^e _r ^{r n} _i ⁿ _i ⁿ g u u u u u m_t ^/ m_t ^/a o p ^a n ^a dn O h ^m s ^o s _i ⁿs ^o s o o o o g u um ^b _t ^y s m e _t ^s s _r ^{2 s} ^o _r _t o ^{2 C} _i ^g u s ^{i o} _t o^{i + h > u} a ,_t u g ^u a r u g _: O O 2 2 _i ^{n o o} g ^{C C H H} w s sS ⁱ _{r l} ^y c _t c _t n h _s p ^a _t a a a e ^r a a e ^{r i} _{f i} ^t u B l^{e w} o ^{H o} B _r o _i ⁿ oⁱ _t oⁱ _t oⁱ _t oⁱ _{t l} ^o _{l l} ^eo _l ^eo g _i ^{n H} _i ⁿ w a ^r a ^r a ^r a ^r o ^f m me m ^{c r} n oi ^r n oi ^o ₍ u ^P g _i ⁿ _s g _s ^o _l ^l _{, ,}g o n d _t d _{t t} B n k a _i ⁿ a o ^e _r ^e _r ^e _r ^e _r m n nn c m n a n a n a _{" i} ^{u oi} _t ^{2 r} _t ^{2 r} _t a ^{H i} g _s a g k r a g_r ^f u m s u s s u u o _i ^o _i ^o s ss ss _f ^{r s s} ^R e d _l ^a dn ne d n ^t _{c i} ^{n u m} a ^m a ^{o e e e e} _{y r r}- n ^e _r o e y z ^a c n e n ^a c a n n ^e _r ^m _r e n h e e n e _{r r r r r} n ^e _v ^e _v h ⁿ e ^{n f} p_l p_l p_l p_l ^a n nu _r ^m n ^m _r ^t _s o ^t _s o _f o ^e _r h ^e _r yn _{i i i i} n o os _l ^o o^t _{c i} ^{o o} _f ¹ n ^c _l a ¹ n ^c _l a i^{m o} a ^f n ^W, _r ^W, _r us W ^a r us ^{a t} _r ^at _r ^at _r ^at _r eh ^{c c}c ^t _c ^o _r ^a _{t i} ^{m a} _t x oⁱ _{t i} ^m e s s n a a a a ^T O O_i ^s ae a d eh e _i ⁿ eh e a a _i ⁿ h ^O O D ^r _" ^e _r y h w m m w m m R m W ^e _r O e ^{P P P P ,} _r ^{C C} p ^e _r p h T O 3 egna R me^t _s ² y e S g _r n o a ^t _c R ae R₎e ^g _ru P ¹ ⁽ eg d n n a 2 R 4 egn 5 5 a R ⁵ _.0 ⁷ _.0 3 egn 0 0 a R ⁵ _.0 ⁸ _.0m ² ^e _t eg ^s _y n 5 5 a S R ⁴ _.0 ⁸ _. _r 0o^t _cae R₎ _i ⁿa ¹ ^M ₍ e _t gn 0 0s a 1 R ⁴ _. 0 ⁹ _. 0 dn a_l a a es _i ^m _i ^mx s _i ⁿ a e v^{m m} n >_, < i^o n _, t_c ⁱ n ⁱ a s r _l s e ^eo _l ^eo^y _r ^h _t m ma o ^/ _tu ^/ _tum ^b s o om e _{s s}u s S ⁱ _r ^e _l p o _l ^eo_s e m mg o _, m n _,nn c m a _{" i} ^u _i ^o ^s _i ^o ^R- en de _re ^s _r v e^e _r oz vus _r ^m n n n o o_l ^o o^t _i ^{o c c}c _c _i ^s a ^t _c 2 2 e a ^{H H} D ^r _" ^e _r Description Applies to All Embodiments [00109] The disclosure that follows applies to all embodiments previously described in this disclosure. [00110] Hydroformylation can occur in any reaction medium known in the art. For example, a multi-phase reaction medium comprises at least one gas phase and at least one liquid phase. In another embodiment, the reaction medium is a two-phase reaction medium comprising a gas phase and a single liquid phase comprising dissolved catalyst compounds. [00111] The reactants in a hydroformylation process can comprise the following: 1) Hydroformylation comprising CO, H2, olefin; or 2) Hydroformylation comprising CO, H2, alpha-olefin; or 3) Hydroformylation comprising CO, H2, ethylene and/or propylene, or 4) Hydroformylation comprising CO, H2, propylene. [00112] The reactant mole fractions in the gas phase exiting a reaction medium and vessel can be at least one selected from the group: 1) Mole % olefin in gas phase exiting reaction medium and vessel > 0.5 %; 2)Mole % olefin in gas phase exiting reaction medium and vessel > 1%; 3) Mole % olefin in gas phase exiting reaction medium and vessel > 2%; 4) Mole % olefin in gas phase exiting reaction medium and vessel > 3 %; 5) Mole % olefin in gas phase exiting reaction medium and vessel < 70 %; 6) Mole % olefin in gas phase exiting reaction medium and vessel < 65 %; 7) Mole % olefin in gas phase exiting reaction medium and vessel < 60 %; 8) Mole % olefin in gas phase exiting reaction medium and vessel < 55 %; 9) Mole % CO in gas phase exiting reaction medium and vessel > 2 %; 10) Mole % CO in gas phase exiting reaction medium and vessel > 3 %; 11) Mole % CO in gas phase exiting reaction medium and vessel > 4 %; 12) Mole % CO in gas phase exiting reaction medium and vessel > 5 %; 13) Mole % CO in gas phase exiting reaction medium and vessel < 60 %; 14) Mole % CO in gas phase exiting reaction medium and vessel < 50 %; 15) Mole % CO in gas phase exiting reaction medium and vessel < 40 %; 16) Mole % CO in gas phase exiting reaction medium and vessel < 30 %; 17) Mole % H₂ in gas phase exiting reaction medium and vessel > 4 %; 18) Mole % H₂ in gas phase exiting reaction medium and vessel > 6 %; 19) Mole % H2 in gas phase exiting reaction medium and vessel > 8 %; 20) Mole % H2 in gas phase exiting reaction medium and vessel > 10 %; 21) Mole % H2 in gas phase exiting reaction medium and vessel < 70 %; 22) Mole % H2 in gas phase exiting reaction medium and vessel < 60 %; 23) Mole % H2 in gas phase exiting reaction medium and vessel < 50 %; and 24) Mole % H₂ in gas phase exiting reaction medium and vessel < 40 %. [00113] These ranges exhibit great variation owing to large differences in reactant solubilities in liquid phase of reaction medium and due to a large number of options for catalysis and for resulting, associated preferences for pressure and temperature of reaction medium. These are all “wet basis values” before removal of vapors of aldehydes and other compounds. The sum of actual gas phase composition ex reaction medium and vessel cannot exceed 100%. The gas phase also comprises inert compounds as disclosed herein. [00114] Catalyst for hydroformylation can be any that is known in the art. Catalyst compositions can be at least one selected from the group: 1) transition metal compounds; 2) soluble transition metal compounds; 3) soluble Rhodium & Cobalt compounds; and 4) soluble rhodium compounds. [00115] The temperature of the first, main reaction medium can be in at least one of the following ranges: 1) > 60°C; 2) > 70°C; 3) > 75°C; 4) > 80°C; 5) < 140°C; 6) < 130°C; 7) < 125°C; and 8) < 120°C. [00116] The pressure of the first, main reaction medium can be in at least one of the following ranges: 1) > 600 kPa ( 87 psia); 2) > 1,000 kPa (145 psia); _{3) > 1,400 kPa (203 psia); 4) > 1,800 kPa (261 psia); 5) < 6,000 kPa (870 psia);} _{6) < 4,000 kPa (580 psia); 7) < 3,200 kPa (464 psia); and 8) < 2,400 kPa(348} psia). [00117] When making exclusively propionaldehyde, the preferred ratio of “stripping gas flow” (defined as all of the gas phase portion exiting a first main reaction medium that is not specifically propionaldehyde vapor) compared to product vapor flow exiting said reaction medium and vessel (the exiting gas phase portion that specifically is propionaldehyde vapor), utilizing any reaction chemistry, any reaction vessel type and any Process Flow Diagram (PFD): 1) > 6:1 molar; 2) > 7:1 molar; 3) > 8:1 molar; 4) > 9:1 molar; 5) < 20:1 molar; 6) < 18:1 molar; 7) < 16:1 molar; and 8) < 14:1 molar. [00118] Using large flows of gas through a first main reaction medium and vessel ensures that substantially all of formed aldehyde is promptly removed from the liquid phase of said reaction medium via the exiting flow of gas phase, thereby minimizing and even substantially eliminating the need for removing a portion of the liquid phase of said reaction medium from said main reaction medium and vessel for the purpose of separating the aldehyde product, e.g., in a liquid phase stripping operation comprising pressure reduction and/or heating operations. [00119] Thus, one embodiment of the invention provides that the portion of formed aldehyde product removed by gas stripping from a first main reaction medium and vessel is the preponderance of the net removal of aldehyde product from said medium and vessel. The ratio of aldehyde removed via gas outlet to net aldehyde removed from reaction medium can be > 80; ratio of aldehyde removed via gas outlet to net aldehyde removed from reaction medium > 90; ratio of aldehyde removed via gas outlet to net aldehyde removed from reaction medium > 100; and ratio of aldehyde removed via gas outlet to net aldehyde removed from reaction medium > 110. [00120] Whenever operating with gas-only product stripping from liquid phase of reaction medium, actual removal rates of aldehyde in overhead gas will perforce be ≥ 100% and will typically be > 110%. This is because some portion of crude condensed aldehyde liquid formed from reaction vessel outlet gas is often refluxed into the demister near top of reaction vessel and because additional crude condensed liquid aldehyde is flashed into the syngas and returned to a reaction medium and vessel via the syngas scrubber. [00121] Also, a reaction medium liquid phase can be removed from a reaction medium and vessel and then be returned without significantly separating aldehyde from said liquid phase. An important example is in provision of an external liquid cooling system, in contrast to placing cooling surfaces within the reaction medium and vessel, in order to satisfy the energy balance. [00122] The net aldehyde removed from a reaction medium and vessel is defined herein as the mass balance sum of all aldehyde flows out of said reaction medium and vessel minus the sum of all aldehyde flows into said reaction medium and vessel averaged over at least about 1 minute. [00123] Over sustained periods of time, the net aldehyde removal rate from a reaction medium and vessel must approximately equal the net chemical formation of aldehyde within said reaction medium and vessel in order to limit sustained accumulation or depletion of mass within said medium and vessel. However, small yield losses of aldehyde occur within said reaction medium and vessel whereby formed aldehyde is converted chemically into formed heavy byproducts. [00124] By using the disclosed, surprisingly large and inventive gas flow ratios, liquid flows from a reaction medium and vessel for the purpose of separating produced aldehyde from catalyst components are minimized and essentially eliminated, subject to the longer-term degradation of catalyst ligand and/or to slow accumulation or depletion of heavy byproducts within said reaction medium and vessel. [00125] However, an external liquid-cooling flow loop may optionally be used for liquid phase energy balance purposes, although a simple liquid-flow cooling loop does not provide a net removal of produced aldehyde, heavy byproducts, and Rh-ligand from said reaction medium and vessel. [00126] In addition, when reaction chemistry is well optimized to minimize the production of heavy byproducts according to the disclosures herein, the removal rate of said byproducts via gas phase exiting a reaction medium can also be matched closely to the production rate of said byproducts, notwithstanding the much lower volatility of said heavy byproducts compared to the volatility of aldehyde products. [00127] Accordingly, the need for continuous removal of the liquid phase of a reaction medium for the purpose of liquid phase mass balance purposes is largely eliminated. Duration of net removing liquid phase of first main reaction medium can be < 12 hours/week, < 8 hours/week, < 4 hours/week, and < 2 hours/week . Frequency of net removing liquid phase of first main reaction medium can range from < 4 events/week, < 3 events/week, < 2 events/week, and < 1 events/week. Mass liquid phase net removed / mass aldehyde net removed from first main reaction medium can be < 4%, < 3%, < 2% and < 1%. [00128] Thus, after initial formulation of a first main reaction medium, it is preferred that only small additional net feed amounts of Rh and/or ligand are added into said first main reaction medium and vessel, and then only infrequently. Mass additional Rh-ligand fed / mass Rh-ligand already in first main reaction medium can be < 8% per week, < 4% per week, < 2% per week, and < 1% per week. Duration of additional feeding Rh and/or ligand to first main reaction medium can be < 12 hours/week, < 8 hours/week, < 4 hours/week, and < 2 hours/week. Frequency of additional feeding Rh and/or ligand to first main reaction medium can be < 4 events/week; < 3 events/week, < 2 events/week and < 1 events/week. [00129] An effective control for reducing the chemical degradation rate of ligand within a reaction medium and vessel is to reduce the operating temperature within a reaction medium and vessel, although corresponding changes are needed for reactant gas composition and/or pressure and/or catalysis within said reaction medium and vessel in order to restore hydroformylation reaction rate to desired target. [00130] Effective controls for reducing the production rate of said byproducts within a reaction medium and vessel are reducing the mole fraction of aldehyde in the liquid phase of said reaction medium and/or reducing the temperature of said medium; and both reductions are facilitated by the surprisingly larger and inventive gas flow ratios disclosed herein. [00131] In addition, it is preferred to operate with relatively large formation rates of aldehyde product per unit mass of liquid phase of reaction medium. Because the total formation rate of heavy byproducts is often about linear with the total mass of said liquid phase, with other factors about constant, increasing the formation rate of aldehyde per unit mass of said liquid phase decreases the formation ratio of heavy byproducts versus aldehyde product. This directly increases yield to aldehyde product; and it also reduces the ratio of less volatile heavy byproducts to more volatile aldehyde products which is important when providing for efficient removal of both simultaneously using a gas phase exiting a first reaction means and medium. The formation rate of aldehyde can be > 12 g-mole/hour/kg of liquid phase of reaction medium; > 14 g-mole/hour/kg of liquid phase of reaction medium; > 16 g-mole/hour/kg of liquid phase of reaction medium; and > 18 g-mole/hour/ kg of liquid phase of reaction medium. [00132] In order to provide such preferred formation rates of aldehyde per unit liquid, it is necessary that all gas-to-liquid mass transfer rates and all chemical reaction rate balances are aptly provided. The hydroformylation rate is achieved by using apt combinations of the catalyst, gas phase concentrations of reactants, pressure, and temperature within the reaction medium as disclosed herein. The mass transfer rate is provided by apt combination of gas flows and reactor vessel geometry as disclosed herein, optionally with mechanical agitation. [00133] The large gas flow ratios, large superficial gas velocities, and large gas phase holdups within reaction medium as disclosed herein are particularly useful for providing a large rate coefficient for gas-to-liquid mass transfer dissolution for olefin and, more especially, for sparingly soluble CO and H2. The mass transfer rate coefficient is often referred to in the art as the product of a mass transfer liquid film coefficient (kL) and a bubble surface area (a), i.e., (kL)*(a) or just kLa. [00134] For hydroformylation reactions, it is particularly important to provide a relatively large kLa and mass transfer capability distributed relatively uniformly throughout said reaction medium, because the sparing equilibrium solubilities of CO and H2 limit both the mass transfer driving force and also the maximum stored capacitance of these reactants within the liquid phase. [00135] A BCR as defined herein is a preferred means for providing both a relatively large value of kLa per unit energy expenditure and for providing this kLa relatively uniformly throughout a reaction medium. [00136] When the formation rate of heavy byproducts is sufficiently suppressed with respect to the formation rate of aldehyde product, the concentration of heavy byproducts within a reaction medium may even begin to decline steadily, tending toward an undesirable loss of volume and mass of reaction medium. One method for sustaining a selected mass and volume inventory of liquid phase within a reaction medium and vessel is to increase the mole fraction of aldehyde in the liquid phase until the formation rate of heavy byproducts is increased back into balance with their removal rate from said reaction medium. However, this is undesirable from the standpoint of reactant yields, and it also causes a greater liquid loading for the recycle gas cooling condenser system disclosed elsewhere herein. [00137] In fact, it may be difficult to match simultaneously the mass balances for both formed aldehydes and formed heavy byproduct in said reaction vessel to their optimized conditions, particularly since it is generally undesirable to increase intentionally the formation rate of said heavy byproducts. [00138] Accordingly, it is preferred to reconcile any excess of net removal of heavy byproducts from a reaction medium and vessel by providing a sustained or intermittent feeding of at least one chemically compatible heavy compound into said reaction medium and vessel. One preferred embodiment is to provide a flow of heavy bottoms byproducts from an aldehyde distillation process, e.g., from the liquid bottoms of a distillation column being used to purify HPr and/or HBu. Another preferred embodiment is to provide a flow of compounds that can be produced by aldol type condensation reactions of aldehydes, e.g., Texanol and IBIB. Yet another preferred embodiment is to provide an initial charge and/or infrequent re-charge of at least one particularly non-volatile and chemically compatible heavy compound, e.g., DOP and DOPT. [00139] It is preferred to operate a first main reaction medium using the following mole fractions of formed aldehyde product in the liquid phase of said reaction medium. Such concentrations provide a useful balancing of the net removal rate of formed aldehyde using the disclosed preferred ranges of reaction pressure, reaction temperature, and stripping gas flow ratios. A larger liquid phase concentration of aldehyde in said reaction medium facilitates the net removal of aldehyde by gas stripping, but said larger concentration adversely increases the formation rate of heavy byproducts thus increasing reactant yield loss. Conversely, a smaller liquid phase concentration of aldehyde usefully suppresses the formation rate of heavy byproducts, but said smaller concentration adversely suppresses the net removal of aldehyde by gas stripping and thus requires excessive flow rates of recycle gas. The concentration of aldehyde in liquid phase of reaction medium can be < 90 mole%; < 85 mole%; < 80 mole%; < 75 mole%. The concentration of aldehyde in liquid phase of reaction medium can be > 20 mole%; > 25 mole%; > 30 mole% and > 35 mole%. [00140] The balance of the liquid phase composition comprises catalyst compounds, heavy byproducts, compatible heavy compounds, and dissolved gases, both reactants and inerts. [00141] As used herein, inerts, inert gases, inert liquids, inert compounds, and non-reactive inerts are defined as all compounds that are inert with respect to a hydroformylation reaction and also with respect to reactions with catalysis compounds. [00142] Compounds that are not inerts, as used herein, comprise CO, H2, olefins, and catalysis poisons. [00143] It is preferred to recover at least a portion of crude aldehyde liquid product by cooling at least a portion of a gas stream exiting a first main reaction medium and vessel to form at least a portion of a crude aldehyde liquid condensate; and it is preferred to separate at least a portion of said condensate from said cooled gas. This is an economically efficient way to separate a crude aldehyde product from Rh-ligand within a reaction medium and vessel, because the energy needed for said separation is largely provided by the exothermic heat of the aldehyde formation reaction itself. [00144] Accordingly, it is preferred to treat a large fraction of said gas stream exiting said reaction medium and vessel in said cooling condenser in order to increase the amount of condensed crude aldehyde liquid separated from said reaction medium. Fraction of gas exiting said reaction medium and vessel fed into a gas cooling condenser can be > 70 mole%, > 80 mole%, and > 90 mole%. The fraction of gas exiting said reaction medium and vessel fed into a gas cooling condenser can be 100 mole%. [00145] It is preferred to use the following pressures and temperatures to facilitate said cooling, condensing and separating of said crude aldehyde liquid condensate from said cooled gas stream. These are preferred based on inherent aspects of the hydroformylation chemistry and thermodynamics coupled with ambient pressure and temperature considerations coupled with mechanical and operating cost factors. 1) Condensing temperature of aldehyde from first main reaction medium and vessel outlet gas can be < 60°C; < 50°C; < 45°C and < 40°C. Condensing temperature of aldehyde from said outlet gas can be > 0°C; > 10°C; > 20°C; and > 25°C . The condensing pressure of aldehyde from said outlet gas can be > 600 kPa ( 87 psia); > 1,000 kPa (145 psia); > 1,400 kPa ( 203 psia) and > 1,800 kPa ( 261 psia). [00146] Various mechanical means for providing said cooling and condensing are known in the art. Those comprising heat exchange surfaces and employing a cooling utility fluid are preferred. Preferred materials for said heat exchange surfaces are disclosed elsewhere herein. Preferred utility fluids for a gas cooling condenser comprise the following. 1) All types of fluids; 2) Liquid fluids; 3) Water of all purities; 4) Ambient and cooling tower water, e.g. oceans, lakes, streams, and evaporatively concentrated cooling tower water, optionally with various corrosion and fouling inhibitors as known in the art. [00147] When an ambient or cooling tower water is used as said cooling utility fluid, it is preferred to use selected ranges for operation of said utility fluid. This is in order to reduce fouling and corrosion of the utility side. Ambient and cooling tower water flow velocity in cooling surface heat exchange means can be > 1.0 m/s; > 1.5 m/s; > 2.0 m/s; and > 2.5 m/s. Ambient and cooling tower water temperature rise in cooling surface heat exchange means can be < 40°K; < 30°K; < 20°; and < 10°K. [00148] When using the ranges of conditions for utility cooling fluids, the controllable range for the amount of cooling and condensing provided by said cooling condenser of fixed size is inherently limited unless gas flow through said cooling condenser is limited. Suitable means for adjusting said gas flow rate in order to manage the removal of heat from a reaction medium and vessel comprise partial bypassing of at least a portion of gas flow around said cooling condenser; substantially isenthalpic flow throttling pressure drop means, e.g., such as flow control valves; and substantially isentropic inlet guide vanes on a radial centrifugal compression means. These are listed in order of increasing preference. [00149] It is preferred to separate at least a portion of said crude liquid aldehyde condensate from at least a portion of the remaining uncondensed cooled gas phase exiting said cooling condenser. Mechanical separating means are preferred. Suitable mechanical separation means are known in the art and comprise gravity separation within said cooling condenser, gravity separation within a separate downstream vessel, e.g. a knock-out tank, a gas- liquid separation cyclone, and a mechanical demister separating means. For very large flows of condensed crude aldehyde liquid, e.g. at typical commercial scales of aldehyde production, a preferred mechanical separating means comprises a liquid knock-out tank, optionally with internal demister means, located closely downstream after said cooling condenser. [00150] In order to maximize the amount of said crude liquid aldehyde that is separated from said cooled gas phase, it is preferred that the separating pressure be as great as possible relative to the smallest pressure of gas exiting said reaction medium and vessel. Smallest pressure in said reaction medium & vessel - separating pressure of said liquid can be < 160 kPa (23 psia), < 120 kPa (17 psia), < 80 kPa (12 psia), and < 40 kPa (6 psia). Cooling liquid phase of reaction medium by contact with heat exchange surfaces [00151] A hydroformylation reaction is highly exothermic, and a hydroformylation reaction medium must by cooled by some means. Preferred means comprise utilizing the enthalpy difference between mass flowing into and out of said reaction medium and vessel and/or contacting at least a portion of said medium and/or at least a portion of the liquid phase of said medium with heat exchange surfaces employing a cooling utility fluid. [00152] Preferred materials for heat exchange surfaces, conduits, vessels and other mechanical means of the inventive process and apparatus comprise: 1) Any metal material fabricated as plate, tubing, piping, wire, mesh, or any other shape. 2) Stainless steels of all grades and types. 3) All variations of stainless steels known in the art as types 304 and 316. [00153] Said heat exchange surfaces may be disposed within said reaction medium and vessel and/or be situated in contact with at least a portion of a liquid phase of said medium that is external to said medium and vessel, with said external liquid phase flowing out from and then back into said medium and vessel. [00154] In one embodiment of the invention at least a part of said cooling exchange surfaces is disposed directly within said medium and vessel. Said internal surfaces may be of any shape and orientation, with preferences as follows: 1) Surfaces oriented having principally an axial orientation, i.e. wherein principally vertical surface area exceeds principally horizontal surface area thereof; 2) Tubular surfaces with said axial orientation; 3) Axial tubular surfaces grouped in at least one tube bundle; and 4) Multiple bundles of axial tubular surfaces disposed approximately uniformly across horizontal sections of said reaction medium for at least about the following heights: (a) Said surfaces traverse > 30% of height wherein said medium diameter ≥ 0.9*D; (b) Said surfaces traverse > 45% of height wherein said medium diameter ≥ 0.9*D; (c) Said surfaces traverse > 60% of height wherein said medium diameter ≥ 0.9*D; and (d) Said surfaces traverse > 75% of height wherein said medium diameter ≥ 0.9*D. [00155] It is preferred to use selected types of utility cooling fluids in order to reduce the surface area of said exchange surfaces disposed within said reaction medium and to reduce corrosion and fouling of the utility coolant side such as: 1) All types of liquid coolants; 2) An enclosed water system cooled by contact with another utility fluid comprising various ambient water supplies, a cooling tower water, and/or ambient air; 3) Said enclosed water system using substantially de-oxygenated and/or de-ionized water, optionally with chemical inhibitors for fouling and corrosion; and 4) Said enclosed, inhibited water system operating with a minimum system pressure of more than about 110 kPa (16 psi) minimum and a system pressure less than about maximum 1,500 kPa (218 psi). [00156] In another embodiment of the invention, at least a portion of cooling exchange surface is located outside of a reaction vessel by utilizing a flow of a portion of liquid phase of said reaction medium out from and back into said reaction medium and vessel, but without significant separation of said liquid phase components, i.e. not significantly separating formed aldehyde from Rh and/or heavy byproducts. [00157] For external cooling of liquid phase of reaction medium, it is preferred to provide at least one liquid withdrawal from said reaction medium and vessel significantly elevated above the lowest elevation at which the diameter of said reaction medium > 0.2*D.1) Elevation liquid phase exiting reaction medium & vessel above lowest elevation with 0.2*D > 2 m 2) Elevation liquid phase exiting reaction medium & vessel above lowest elevation with 0.2*D > 3 m 3) Elevation liquid phase exiting reaction medium & vessel above lowest elevation with 0.2*D > 4 m 4) Elevation liquid phase exiting reaction medium & vessel above lowest elevation with 0.2*D > 6 m. [00158] For external cooling of liquid phase of reaction medium, it is preferred to provide at least one liquid withdrawal from said reaction medium and vessel significantly elevated above the lowest elevation at which the diameter of said reaction medium >0.2*D). [00159] Various means of gas-liquid separation are known in the art and disclosed herein. Gravitational separation by a density difference between a liquid phase and a gas phase is preferred. In said gravitational embodiment, it is preferred that the vertical, downwards superficial velocity of liquid phase be slowed sufficiently crossing at least one imaginary horizontal plane cutting across a flow conduit or vessel conducting said externally cooled liquid phase. Said slowed downwards superficial liquid velocity promotes preferred ranges of deaeration by gravity. 1) Superficial downwards time-averaged velocity of externally cooled liquid can be < 0.5 m/s (1.6 ft/s) < 0.3 m/s (1.0 ft/s); < 0.2 m/s (0.7 ft/s); and < 0.1 m/s (0.3 ft/s). [00160] In one embodiment, at least a portion of said deaeration of cooling liquid flow occurs outside of said reaction vessel. In said external deaeration embodiment, it is preferred that at least a portion of said externally separated gas phase reenters said reaction vessel at an elevation above the lowest elevation of its withdrawal in order to promote a pressure differential to support the external circulation of liquid flow. Gas reentry can be above withdrawal elevation > 1 m; > 2 m; > 3 m; and > 4 m. [00161] In one embodiment of external deaeration of cooling liquid flow, it is preferred that said externally separated portion of gas phase reenters said reaction medium and vessel at a level below H, the elevation of the top of said reaction medium. [00162] In another embodiment of external deaeration of cooling liquid flow, it is preferred that said externally separated portion of gas phase reenters said reaction vessel at an elevation located above H and within the gas ullage elevations of said vessel. Such locations can be: 1) below exiting height of preponderance of gas phase from reaction vessel; 2) below exiting height of preponderance of gas phase from reaction vessel by > 0.4 m; 3) below exiting height of preponderance of gas phase from reaction vessel by > 0.5*D and 4) below at least one mechanical demister means located within said reaction vessel. [00163] When said external liquid cooling surfaces are used, it is preferred to limit the dissipation of energy due to the external flow of liquid phase within the following ranges of static pressure loss: 1) external cooling liquid flow loop pressure loss, maximum - minimum pressure, < 700 kPa (102 psia); 2) external cooling liquid flow loop pressure loss, maximum - minimum pressure, < 525 kPa ( 76 psia); 3) external cooling liquid flow loop pressure loss, maximum - minimum pressure, < 350 kPa ( 51 psia); and 4) external cooling liquid flow loop pressure loss, maximum - minimum pressure, < 175 kPa ( 25 psia). [00164] The external flow cooling flow of said deaerated liquid may be pumped by a mechanical means, with said mechanical pumping means preferably disposed after and with at least a portion thereof at a lower elevation than said deaeration means. Optionally, said external cooling liquid flow may occur without a mechanical pumping means and thus be owing to gravitational force working with the density difference between lighter aerated reaction medium and denser deaerated liquid phase thereof. [00165] When providing external cooling liquid flow without a mechanical pumping means, it is preferred that the highest elevation at which said deaerated liquid is formed is within selected ranges above at least one elevation at which said deaerated cooling liquid flow reenters said reaction medium and vessel at the following locations: 1) Elevation liquid phase exiting reaction medium - cooled liquid reentering reaction medium > 3 m; 2) Elevation liquid phase exiting reaction medium - cooled liquid reentering reaction medium > 5 m; 3) Elevation liquid phase exiting reaction medium - cooled liquid reentering reaction medium > 7 m; and 4) Elevation liquid phase exiting reaction medium - cooled liquid reentering reaction medium > 9 m. [00166] It is preferred for at least a portion of said cooling liquid flow to reenter said reaction medium and vessel at a relatively low elevation in the following locations: 1) Elevation cooled liquid reentering reaction medium above bottom of reaction medium < D m; 2) Elevation cooled liquid reentering reaction medium above bottom of reaction medium < D/2 m; 3) Elevation cooled liquid reentering reaction medium above bottom of reaction medium < D/3 m; and 4) Elevation cooled liquid reentering reaction medium above bottom of reaction medium < D/4 m. [00167] In one embodiment, it is preferred for at least a portion of said cooling liquid flow to reenter said reaction medium below the elevation at which the diameter of said reaction medium and vessel first exceeds 0.5*D and above the elevation where the majority of recycle gas first enters said reaction medium. Said return elevation enables recycle compressor energy to provide additional gas-lift work to promote availability of a greater pressure differential for gravity circulation of external cooling liquid at the following locations: 1) Elevation cooled liquid reentering below (lowest elevation where reaction medium > 0.5*D) > 0.25 m; 2) Elevation cooled liquid reentering below (lowest elevation where reaction medium > 0.5*D) > 0.50 m; 3) Elevation cooled liquid reentering below (lowest elevation where reaction medium > 0.5*D) > 0.75 m; 4) Elevation cooled liquid reentering below (lowest elevation where reaction medium > 0.5*D) > 1.0 m; 5) Elevation cooled liquid reentering below (lowest elevation where reaction medium > 0.5*D) < 8 m; 6) Elevation cooled liquid reentering below (lowest elevation where reaction medium > 0.5*D) < 6 m; 7) Elevation cooled liquid reentering below (lowest elevation where reaction medium > 0.5*D) < 4 m; and 8) Elevation cooled liquid reentering below (lowest elevation where reaction medium > 0.5*D) < 3 m. [00168] Such a mechanical arrangement may be provided by various means comprising providing a single reaction vessel having a cylindrical cross section of reduced diameter near its lower elevations, providing at least one smaller diameter reaction vessel closely connected and more preferably below a larger diameter first main reactor vessel, and/or providing a suitable section of smaller diameter conduit closely connected and more preferably below a larger diameter first main reactor vessel. [00169] Even with the preferred embodiments herein, the differential pressure available for providing a gravitationally driven external flow of cooling liquid remains relatively small, in comparison to providing a mechanical pumping means; and it is important to limit the flowing pressure drop of said cooling liquid. In this regard, it is preferred to provide selected ranges for the superficial flow velocity of said liquid phase in at least one conduit carrying said external cooling liquid flow, and more preferably for the preponderance of all external cooling liquid flow. The superficial velocity of liquid phase of reaction medium in external cooling flow can be < 2.5 m/s; < 1.5 m/s; and < 1.2 m/s. The superficial velocity of liquid phase of reaction medium in external cooling flow can be > 0.5 m/s; > 0.6 m/s; > 0.7 m/s; and > 0.8 m/s. [00170] Said velocities of cooling liquid provide a balancing between the flow energy requirement and the mass of liquid phase of reaction medium and/or mass of expensive Rh deployed in a deaerated and cooled condition, i.e. relatively inactive chemically, outside of a first main reaction medium and vessel. [00171] Various mechanical means for providing cooling of said external cooling liquid flow are known in the art. Those comprising heat exchange surfaces and employing a cooling utility fluid are preferred. Preferred materials for said surfaces are disclosed herein. Preferred utility fluids and operating ranges for said utility fluids are disclosed herein. [00172] An advantage of external cooling liquid flow is that a less costly utility fluid operation may be more practicable for external exchange surface than for heat exchange surface located within a reaction vessel. [00173] Preferred ranges of operating conditions for water utility fluids, as disclosed herein, cause limitations for changing the heat duty of a heat exchange means when using unregulated flow of external cooling liquid. Accordingly, it is preferred to provide a greater range for regulating the amount of cooling by providing regulation of the flow rate of external liquid cooling. Preferred means for regulating said flow comprise variable speed pumping means and/or substantially isenthalpic pressure and flow throttling valve means. When using gravitational flow of external cooling liquid without a mechanical pumping means, it is preferred to use at least 2 of said throttling means in parallel on said flow of external cooling liquid in order to provide a greater range of finer control of the flow rate of said external liquid flow and hence on the cooling duty of said heat exchange means. [00174] It is preferred that the radial and azimuthal position disclosures for feeding recycle gas and makeup reactant feeds are applied also for radial and azimuthal distributions of liquid reentering a reaction medium and vessel from an external liquid cooling flow loop. Preferred condenser cooling duty, with any reaction vessel type with any Rh- ligand with any PFD [00175] Using a relatively larger flow of gas through a first main reaction medium and vessel followed by cooling a preferred preponderance of gas exiting said medium and vessel in a gas cooling condenser results in a relatively greater heat duty for said gas cooling condenser and a resultingly lesser heat duty for cooling the liquid phase of said reaction medium by direct contact with heat exchange surfaces. [00176] It is surprisingly beneficial to shift the process cooling duty toward said cooling condenser in this manner notwithstanding that a more typical preference is to place surface heat exchange cooling duty on a liquid phase rather than a gas phase of similarly constituted fluids. Surprising benefits of this aspect of the present invention comprise the following reducing or eliminating the need to place heat exchange surfaces inside the expensive multiphase reaction vessel, which location excludes reaction medium thereby enlarging reaction vessel size and often increasing various downtime and maintenance costs, and/or reducing or eliminating the flow rate and volume of liquid phase that is circulated out of and back into said reaction vessel through an external cooling loop comprising liquid phase heat exchange surfaces, which flow loop increases the working inventory of Rh-ligand and which inherently shifts balances of chemical reaction rates between products and byproducts, and possibly between product isomers. [00177] In another embodiment of the invention, the gas flow ratio provided to a first main reaction medium and vessel is sufficiently large such that the recycle gas cooling condenser operation satisfies the majority of the energy balance for said reaction medium. This embodiment greatly limits the capital and operating costs associated with providing heat exchange surfaces within said reaction vessel and/or providing heat exchange surfaces operating on a liquid flow cooling loop exiting and returning directly to said reaction medium and vessel. The recycle gas condenser heat duty/aldehyde formation heat duty can be > 50%; > 55% ; > 65%; and > 75%. [00178] The heat duty of the recycle gas condenser is defined herein as the enthalpy difference between all process flow entering said condenser and all process flow exiting said condenser. [00179] The aldehyde formation heat duty is defined herein as the enthalpy difference between equimolar amounts of gaseous olefin, CO, plus H2 and the resulting produced aldehyde vapor, all evaluated at the temperature of said reaction medium and the pressure at the highest elevation of said reaction medium and notwithstanding that hydroformylation reaction occurs in the liquid phase. [00180] In another embodiment of the invention, the gas flow ratio provided to a first main reaction medium and vessel is sufficiently large to support the preponderance of the cooling duty for the entire synthesis operation providing a crude aldehyde liquid. Besides the various exothermic heats of reaction within said reaction medium and vessel, an aldehyde synthesis process as disclosed herein often comprises shaft driven energy inputs, e.g., comprising recycle gas compressor, liquid pumps, and perhaps mechanical agitation of a reaction medium; and there are often heat input exchange surface in a synthesis process as well, e.g., comprising heat input to control the temperature in a syngas scrubber unit operation wherein condensed crude liquid aldehyde is contacted with fresh makeup synthesis gas. Recycle gas condenser heat duty/aldehyde formation heat duty can be > 96%; > 100%; > 102% and > 104%. [00181] A particular utility of this embodiment of the invention is virtually or completely eliminating the need for heat exchange surfaces to contact the liquid phase of said reaction medium, either within said reaction vessel or in an external liquid flow cooling loop. [00182] Said syngas scrubber is optionally provided for the purposes of absorbing at least one impurity compound from a fresh makeup synthesis gas supply before said syngas is first provided to a reaction medium and/or for reducing the amount of olefin dissolved in crude liquid aldehyde exiting said syngas scrubber. Recycle gas total compression duty, and dP across reaction medium and vessel [00183] After separating preferred amounts of crude liquid condensate from a cooled gas phase exiting said cooling condenser, it is preferred to send a very large fraction of said cooled, separated gas phase back into said first main reaction medium and vessel to achieve greater cumulative conversions of fresh makeup reactants. This gas return to an upstream condition within a process requires a gas compression duty. [00184] It is preferred to provide said gas compression duty using a mechanical compression means. Gas fed into and exiting from said compression means is herein called recycle gas because most of this compressed flow is typically recycled back into said reaction medium and vessel. A process assembly comprising recycle gas flow conduits, heat exchange means, liquid separation means, flow control means, and compression means is herein called a recycle gas flow loop; and said compression means is herein called a recycle gas compressor. [00185] It is preferred that said recycle gas compressor comprises a rotating shaft and operates with good thermodynamic efficiency relative to a minimum requirement for an ideal isentropic increase of pressure. Thermodynamic efficiency of compression relative to ideal isentropic compression can be > 70%; > 75%; > 80% and > 85%. [00186] A preferred mechanical compression means is all types of single- stage, radial flow centrifugal compressors, as known in the art. [00187] A preferred embodiment of said compressors is one driven by an electrical motor operating directly at the speed of a 2-pole induction or synchronous electrical motor, e.g., about 3,600 rpm on 60 Hz power, or about 3,000 rpm on 50 Hz power. This design provides a good economy balancing compressor flow and head duties preferred for commercial-scale production of aldehydes without requiring too large of a compressor housing and impeller on the one hand or too great of a rotational speed on the other. [00188] Said compressors are typically robust mechanically and often do not require complete demisting removal of condensate exiting the recycle gas cooling condenser before uncondensed recycle gas enters said compressor suction. However, it is preferred that a recycle gas flow entering said recycle compressor suction does not comprise slugs of liquid and that the mass fraction of liquid mist is in preferred ranges at the compressor suction. Mass of mist liquid/total mass of suction stream can be < 8%; < 2% and < 1%. [00189] An embodiment of said compressors preferred for commercial-scale production of aldehydes further comprises inlet guide vanes for control of flow capacity versus compression head. Optionally, additional flow capacity control for the recycle gas loop may be provided by a variable speed electrical motor on said compressor and/or by a pressure reduction means located near said compressor suction or discharge. [00190] It is preferred to optimize the mechanical design of a first main reaction vessel and recycle gas flow loop such that the pressure increase required across said recycle gas compressor is within selected ranges. For the inventively large gas flow rates per unit aldehyde production as disclosed herein, these compression ranges provide a useful economy for an efficient design of a first main reaction medium combined with a recycle gas flow loop. Capital and energy requirements of the recycle gas flow loop comprise static pressure losses due to gas flow related to aeration and mixing of said reaction medium, de-entraining Rh-ligand from the gas phase exiting said reaction medium and vessel, flow conduits for recycle gas, cooling and condensing means for recycle gas, separating means for condensed crude aldehyde liquid, controls to adjust the flow rate of recycle gas as needed, and distributing recycle gas back into said reaction medium and vessel. dP of gas flowing through recycle compressor can be < 550 kPa increase (80 psia); < 450 kPa increase (65 psia); < 350 kPa increase (51 psia); < 250 kPa increase (36 psia). dP of gas flowing through recycle compressor can be > 70 kPa increase (10.2 psia); > 80 kPa increase (11.6 psia); > 90 kPa increase (13.1 psia); and > 100 kPa increase (14.5 psia). [00191] It is preferred to optimize gas flow rates, reaction medium density, and physical dimensions of a first main reaction medium and vessel using disclosure herein such that the gas-flow pressure drops through said reaction medium and vessel are within the following ranges. These ranges provide a useful economy of capital and energy costs needed to control interphase and bulk gradients of composition and temperature within said reaction medium and vessel. dP of gas flowing through first main reaction medium and vessel can be < 160 kPa decrease (23 psia); < 130 kPa decrease (19 psia); < 100 kPa decrease (15 psia); and < 70 kPa decrease (10 psia). dP of gas flowing through first main reaction medium and vessel can be > 15 kPa decrease (2.2 psia); > 20 kPa decrease (2.9 psia); > 25 kPa decrease (3.6 psia) and > 30 kPa decrease (4.4 psia). [00192] In another embodiment, a process is provided to give larger gas flows to provide reduced axial gradients of CO and/or H2 gas phase partial pressure when operating with selected reactor pressures, e.g., making greater n/i-HBu product ratios, with any reaction vessel type with any Rh-ligand with any PFD I.e., employing all other disclosures herein and in all possible combinations and subject to selected boundaries disclosed, which are often useful for producing high ratios of n/i HBu using relatively smaller partial pressures of sparingly soluble CO and/or H2. [00193] When pressure of gas phase exiting reaction medium and vessel is 1) > 1,000 kPa (145 psi); 2) > 1,300 kPa (189 psi); 3) > 1,600 kPa (232 psi); 4) > 1,900 kPa (276 psi); [00194] And when either or both partial pressure of gaseous CO within reaction vessel outlet gas phase 1) < 170 kPa (25 psi); 2) < 140 kPa (20 psi); 3) < 110 kPa (16 psi); 4) < 80 kPa (12 psi); and [00195] partial pressure of gaseous H2 within reaction vessel outlet gas phase is 1) < 340 kPa (49 psi); 2) < 280 kPa (41 psi); 3) < 220 kPa (32 psi); and 4) < 160 kPa (23 psi); [00196] Then, provide a sufficiently large ratio of gas flow to formed aldehyde such that either or both the ratio of (CO partial pressure in gas phase exiting a first main reaction medium and vessel)/(CO partial pressure calculated from the sum of recycle gas plus all feed inlet streams entering said medium and vessel) is one of the following: 1) CO outlet partial pressure /CO inlet partial pressure > 25%; 2) CO outlet partial pressure /CO inlet partial pressure > 30%; 3) CO outlet partial pressure /CO inlet partial pressure > 35%; and 4) CO outlet partial pressure /CO inlet partial pressure > 40%. [00197] The ratio (H₂ partial pressure in gas phase exiting a first main reaction medium and vessel)/(H2 partial pressure calculated from the sum of recycle gas plus all feed inlet streams entering said medium and vessel) can be any of the following: 1) H2 outlet partial pressure /H2 inlet partial pressure > 50%; 2) H2 outlet partial pressure /H2 inlet partial pressure > 55%; 3) H2 outlet partial pressure /H2 inlet partial pressure > 60%; and 4) H2 outlet partial pressure /H2 inlet partial pressure > 65%. [00198] These reduced axial gradients of partial pressure of CO and/or H2 are particularly useful and preferred when operating with greater pressure and when producing greater n/i ratios of HBu within preferred ranges of the following: 1) > 5:1; 2) > 10:1; 3) > 20:1; and 4) > 30:1. [00199] More preferably, employ these increased gas rates and/or these more uniform gas composition ratios, outlet to inlet, along with all other apt disclosures, in a BCR lacking significant mechanical agitation and having preferred physical dimensions and ratios. [00200] Preferred mechanical agitation power input: can be any of the following: 1) < 2 kW/m3 reaction medium and/or vessel volume; 2) < 1 kW/m3 reaction medium and/or vessel volume; 3) < 0.5 kW/m3 reaction medium and/or vessel volume; and 4) < 0.25 kW/m3 reaction medium and/or vessel volume. [00201] Preferred gas work (V * dP) agitation power input can be the following: 1) > 0.6 kW/m3 reaction medium and/or vessel volume; 2) > 0.8 kW/m3 reaction medium and/or vessel volume; 3) > 1.0 kW/m3 reaction medium and/or vessel volume; and 4) > 1.2 kW/m3 reaction medium and/or vessel volume. [00202] Preferred time-averaged superficial gas velocity Ug in a first main reaction medium and vessel, both BCR and CSTR, especially BCR, can be the following: 1) > 0.12 m/s; 2) > 0.16 m/s; 3) > 0.20 m/s; and 4) > 0.24 m/s. [00203] Preferred time-averaged superficial gas velocity Ug in a first main reaction medium and vessel, both BCR and CSTR, especially BCR, can be the following: 1) < 0.8 m/s; 2) < 0.7 m/s; 3) < 0.6 m/s; and 4) < 0.5 m/s. [00204] In some embodiments, preferred gas holdup Eg in a first main reaction medium and vessel, both BCR and CSTR, especially BCR. (too small is too little kLa per unit liquid at Oxo gas dissolution demand) can be: 1) > 20 vol%; 2) > 25 vol%; 3) > 30 vol%; and 4) > 35 vol%. [00205] Preferred gas holdup Eg in a first main reaction medium and vessel, both BCR and CSTR, especially BCR. (too big is too little liquid per vessel cost) can be: 1) < 70 vol%; 2) < 60 vol%; 3) < 55 vol%; and 4) < 50 vol%. Preferred large D, short H, small H/D of reaction vessel, including both BCR and CSTR [00206] Scale matters in multiphase mixing and directly affects the absolute amounts and balances of bulk convective mixing, micro-mixing, interphase mass transport, and mixing power consumption, with and without mechanical agitation means. The scale factors, H, D, and H/D all inherently matter separately and interactively for magnitudes and ratios of axial, radial and azimuthal mixing flow strengths and for interphase mass transfer. [00207] The effect of scale is, however, particularly important in a natural convection system, such as a BCR, because adjustable design aspects of mechanical agitation providing are lacking. [00208] D is defined as the largest horizontal inside diameter of a first main reaction medium and of a first main reaction vessel, excluding vessel piping nozzles. [00209] A value of D that is too small increases gas hold-up Eg to be too large a fraction of vessel volume, reduces end-to end mixing of liquid phase, other factors constant, and may lack commercial importance 1) D > 2.5 m 2) D > 3.0 m 3) D > 3.5 m 4) D > 4.0 m. [00210] A value of D that is too large reduces gas hold up to be too small a fraction of vessel volume reducing gas-to-liquid mass transfer rates and may cause a vessel to be too expensive relative to production capacity and can have the following: 1) D < 12 m; 2) D < 11 m; 3) D < 10 m; and 4) D < 9 m. [00211] H is defined as the vertical height of first main reaction medium having horizontal diameter ≥ 0.2*D. [00212] A value of H that is too tall increases the amount of reaction medium to be too large (too much Rh-ligand and too much adverse chemistry), increases end-to-end mixing times of liquid phase, and increases recycle gas compressor dP beyond efficient values. H can be < 30 m; < 26 m; < 22 m; and < 18 m. [00213] A value of H that is too short pushes D to be very large for a given amount of mother liquor in the vessel thereby driving Ug and Eg to be inefficiently small relative to interphase mass transfer needs; and a too short H may also cause a reaction vessel mechanical design and capital cost to be too expensive relative to production capacity. H can be > 3 m; > 5 m; > 7 m; and > 9 m. [00214] A sufficiently large ratio of H/D within an Oxo main BCR avoids too much vessel cost and too little interphase mass transfer of reactant gases. H/D can be > 1:1; > 1.2:1; > 1.4:1; and > 1.6:1. [00215] A sufficiently small ratio of H/D within an Oxo main BCR avoids too much power consumption at the recycle gas compressor and helps limit the axial gradients of temperature and composition within the BCR, as disclosed elsewhere herein. H/D can be < 9:1; < 7:1; < 5:1; and < 3:1. Preferred de-entrainment of liquid from gas exiting reaction medium and vessel [00216] Several aspects of the invention are directed toward limiting the entrainment of catalyst bearing liquid phase of a first main reaction medium within a flow of gas exiting said reaction medium and vessel. [00217] It is preferred to locate within a first main reaction vessel at least one means whereby slugs, droplets and mists of the liquid phase are removed from the exiting gas phase. This reduces the capital and operating costs compared to placing a separating means in another vessel connected by flow conduits. [00218] One preferred means for separating liquid phase from the exiting gas phase is to allow the action of gravitational force to separate larger slug and droplet portions of the denser liquid phase that are erupting upwards from said reaction medium, propelled by vertical momentum locally near the top of said reaction medium. [00219] The height L of a reaction vessel is defined herein as the height from the lowest elevation at which the inside diameter of the vessel exceeds 0.2*D extending upwards to highest elevation at which the inside diameter of the vessel is reduced below 0.2*D. [00220] It is preferred for the preponderance of gas phase exiting said reaction vessel and flowing through conduits to the recycle gas cooling and compression means to exit quite near the top of said reaction vessel, disclosed as follows: 1) L- gas exiting elevation < 0.4 m; 2) L- gas exiting elevation < 0.3 m; 3) L- gas exiting elevation < 0.2 m; and 4) L- gas exiting elevation < 0.1 m. [00221] It is preferred to limit within selected ranges the mass fraction of liquid phase entrained in a predominantly gas phase flow exiting a first main reaction medium and vessel, even when said liquid phase does not comprise catalyst compounds. Among other benefits, this reduces difficulties with accumulation of liquid slugs in conduits leading to a recycle gas condensing means. Effective means for limiting said entrained liquid fraction comprise vessel ullage height and mechanical demister means, as disclosed herein. Mass of liquid flow entrained in mass of outlet gas flow can be < 4%; < 2%; < 1%; and < 0.5%. [00222] In addition, it is preferred to limit even more strictly the entrainment of catalyst compounds in a predominantly gas phase flow exiting a first main reaction medium and vessel. [00223] The ullage height within a reaction vessel is defined herein as L-H. [00224] The following preferred ranges for L-H allow for a useful control range for operating the level of aerated reaction medium plus also providing height within a first main reaction vessel for various means of separating liquid phase from the exiting gas phase. L-H can be > 1.0 m; L-H > 1.5 m; L-H > 2.0 m; L- H > 2.5 m; L-H < 6.5 m; L-H < 5.5 m; L-H < 4.5 m; and L-H < 3.5 m. [00225] In consideration of preferred ranges of H and D for a reaction medium and of ullage height for a reaction vessel, selected ranges of L and L/D are preferred. L can be > 4 m; L > 8 m; L > 10 m; L > 12 m; L < 32 m; L < 28 m; L < 24 m; and 8) L < 20 m. L/D can be > 1.2; L/D > 1.4; L/D > 1.6; L/D > 1.8; L/D < 10; L/D < 8; L/D < 6; and L/D < 4. [00226] It is preferred to locate at least one mechanical demisting means within the ullage height and near the upper elevation of a reaction vessel. Preferred demisting means comprise demister pads, plate demisters, vane demisters and other means known in the art of gas-liquid separation. [00227] Selected locations are preferred for a demister means within a reaction vessel including: 1) Bottom of demister elevation above H > 0.4 m; 2) Bottom of demister elevation above H > 0.8 m; 3) Bottom of demister elevation above H > 1.2 m; 4) Bottom of demister elevation above H > 1.6 m; 5) Top of demister elevation below outlet of majority of gas phase from reaction vessel > 0.10*D; 6) Top of demister elevation below outlet of majority of gas phase from reaction vessel > 0.15*D; 7) Top of demister elevation below outlet of majority of gas phase from reaction vessel > 0.20*D; and 8) Top of demister elevation below outlet of majority of gas phase from reaction vessel > 0.25*D. [00228] To further maximize the retention of catalyst components within a first main reaction vessel, it is preferred to provide a flow of a compatible cleansing liquid within a reaction vessel and well distributed horizontally above H and more preferably above at least one mechanical demisting means. Many suitable types of liquid distribution means are known within the art comprising spray heads, flooded perforated trays, perforated conduits, and the like. A preferred compatible cleansing liquid is a portion of condensed crude aldehyde liquid removed by condensation from the recycle gas flow, and this is defined herein as aldehyde condensate reflux. [00229] To ensure good liquid wetting and downwards cleansing of the demisting means, the following liquid cleansing ratios are preferred, expressed as a mass fraction of the production rate of aldehyde product within said reaction medium. Liquid cleansing flow / aldehyde production rate can be > 1%; > 2%; > 3%; and > 4%. [00230] There are numerous ways to increase or to decrease the formation rate of crude aldehyde condensate. Methods to increase the flow of crude liquid aldehyde in the recycle gas cooling condenser comprise increasing the flow of recycle gas, reaction temperature, and/or mole fraction of aldehyde in the liquid phase of a reaction medium and by decreasing reaction pressure and/or outlet temperature of a recycle gas cooling condenser. The reverse actions will decrease the formation rate of crude aldehyde condensate. [00231] However, it is not necessary that the formation rate of crude aldehyde condensate in a recycle gas cooling condenser matches precisely with the various process flow uses for said condensate, said uses comprising supply of aldehyde condensate reflux flow to a reaction vessel demister means, re- evaporation of crude aldehyde liquid within a syngas purifying and olefin recovery scrubber means, removal of crude aldehyde product to match aldehyde production rate, and other process needs. An excess of said condensate above other process uses can be returned to said reaction medium and vessel, either above or below the demisting means for the gas exiting said reaction medium and vessel. [00232] However, providing said condensate above the demister will, in addition to supporting retention of Rh-ligand within said reaction medium, also provide a distillation effect retaining a greater portion of higher boiling heavy byproducts within said reaction medium. Retaining too much of the heavy byproducts within reaction medium is undesirable whenever their retention rate exceeds their formation rate over extended periods of time; for then the level of reaction medium will rise until intervention using an undesirable purge of a portion of liquid phase comprising costly Rh-ligand. [00233] Thus, upper limits of reflux flow ratio of aldehyde condensate above a reaction vessel demister means are preferred, expressed as a mass fraction of the production rate of aldehyde product within a first main reaction medium. Aldehyde condensate reflux flow / aldehyde production rate can be < 50%; < 40%; < 30%; and < 20%. Gas and liquid feeding locations into reaction medium and vessel [00234] To provide good utilization of a first main reaction medium and of the compression power used to recycle gas, the following axial locations are preferred for at least a portion recycle gas and/or for at least a portion of fresh makeup reactant supplies. Recycle gas and reactant feeding elevation can be < 0.6*D above bottom elevation reaction medium; < 0.5*D above bottom elevation reaction medium, < 0.4*D above bottom elevation reaction medium, and < 0.3*D above bottom elevation reaction medium. [00235] The bottom elevation of a reaction medium and vessel is defined herein as the lowest elevation where the diameter of said reaction medium and vessel exceeds about 0.2*D. [00236] Often, is useful to comingle and feed a preponderance of fresh makeup reactants along with recycle gas flow; but in other situations, it may be useful to feed at least a portion of fresh makeup reactants separated from the preponderance of recycle gas flow. [00237] In one embodiment of the invention, at least a portion of fresh makeup propylene is fed as a pressurized liquid separated from recycle gas flow. This separation reduces mechanical design difficulties that can occur after mixing a portion of propylene liquid with a portion of recycle gas, said difficulties comprising hydraulic hammer due to liquid slug flow and cryogenic flashing of propylene liquid into recycle gas. [00238] In another embodiment of the invention, at least a portion of the fresh makeup CO and/or H₂ are fed to a first main reaction medium and vessel separated from recycle gas and entering into said reaction medium and vessel above the feeding elevation for a majority of recycle gas. As already noted herein, it is sometimes particularly useful to limit axial gradients of reactant concentrations in the gas phase, especially sparingly soluble CO and H2; and said separated, elevated feeding of a portion of CO and/or H2 is useful in this regard. [00239] For example, providing about one-half of the fresh makeup CO at an elevation of about 0.5*H above the bottom of a reaction medium can reduce the axial gradient in gas phase CO concentration by about one-half without changing the recycle gas flow rate. Of course, this is an approximate analysis before considering the feeding means and mixing methods for dispersing the gaseous CO and before considering the liquid mixing and its limited capacitance for dissolved CO. [00240] Accordingly, the following ranges are disclosed for obtaining additional suppression of concentration gradients of CO and/or H2 whenever the additional mechanical complexity is sufficiently rewarded by improvements in hydroformylation chemistry. [00241] A portion of fresh makeup CO and/or H2 added above feeding elevation of preponderance of recycle gas can include: 1) Fraction of total fresh makeup flow of individual reactant < 90 mole%; 2) Fraction of total fresh makeup flow of individual reactant < 80 mole%; 3) Fraction of total fresh makeup flow of individual reactant < 70 mole%; 4) Fraction of total fresh makeup flow of individual reactant < 60 mole%; 5) Fraction of total fresh makeup flow of individual reactant > 10 mole%; 6) Fraction of total fresh makeup flow of individual reactant > 20 mole%; 7) Fraction of total fresh makeup flow of individual reactant > 30 mole%, and 8) Fraction of total fresh makeup flow of individual reactant > 40 mole%. [00242] Elevation of feeding said separated fresh makeup portion of CO and/or H2 can be 1) Elevation of feeding said gas portion > H/6 above bottom elevation of reaction medium; 2) Elevation of feeding said gas portion > H/5 above bottom elevation of reaction medium; 3) Elevation of feeding said gas portion > H/4 above bottom elevation of reaction medium; and 4) Elevation of feeding said gas portion > H/3 above bottom elevation of reaction medium. [00243] Certain azimuthal and radial locations for feeding makeup reactants and recycle gas are also preferred, in addition to the feeding elevations. [00244] It is preferred to feed recycle gas and makeup reactant gases approximately uniformly distributed azimuthally at each elevation of feeding. [00245] In one preferred embodiment, the radial feeding of recycle gas and/or makeup reactants is provided very near the horizontal center near the bottom of a first main reaction medium and vessel, e.g., through an open conduit entering near the bottom center of said reaction medium. The utility of this embodiment is affected by the absolute scale of a reaction medium and by the disposition of any internals within said reaction medium, e.g. heat exchange surfaces, mechanical agitation means, and flow baffling means. [00246] In another preferred embodiment particularly suitable for BCR embodiments having the preferred short H/D disclosed herein, multiple feeding locations for recycle gas and/or makeup reactants are provided approximately uniformly radially over substantially the entire cross section of a first reaction medium and vessel at the selected feeding elevation, e.g., using a substantially horizontal perforated plate gas sparger fitted closely to the inside wall of a reaction vessel, as is known in the art. [00247] In another preferred embodiment, gas sparger conduits are used to feed the recycle gas and/or makeup reactants approximately uniformly radially at the selected feeding elevations. Such spargers may comprise one or more piping rings or conduit polygons; or they may comprise radial conduit spokes. [00248] Still other conduit network and/or distribution plenum configurations may be used. [00249] To control the compression energy required for the disclosed, inventively greater flows of recycle gas, in balance with the disclosed distribution locations of recycle gas, there are preferred ranges of flowing pressure drops for recycle-gas feeding mechanical means. These pressure drops consider the chaotic pressure fluctuations of the 2-phase gas-liquid system near the gas distribution means coupled with the compression energy expended coupled with mixing factors for optimizing said reaction medium. The gas feeding means can have a dP > 3 kPa (0.4 psia); dP > 6 kPa (0.9 psia); dP > 9 kPa (1.3 psia); dP > 12 kPa (1.7 psia); dP < 160 kPa (23 psia); dP < 120 kPa (17 psia); dP < 80 kPa (12 psia); and dP < 40 kPa ( 6 psia). [00250] Said dP is measured from where said gas feed first traverses the outer boundaries of said reaction vessel up to the point at which said gas contacts reaction medium substantially continuously. Axial thermal gradients within reaction medium [00251] Thermal gradients within a reaction medium and vessel affect the balance between multiple chemical reactions comprising the intended formation of aldehyde products and the isomer ratios thereof, adverse byproduct reactions, and catalyst ligand degradation reactions. Thermal gradients arise from the widely distributed aspect of the exothermic chemical reactions coupled with the more localized placement of gas and liquid feeding locations and heat exchange surfaces. [00252] It is preferred to limit the axial thermal gradient of said reaction medium within the following ranges: 1) dT < 12°K; 2) dT < 9°K; 3) dT < 6°K; and 4) dT < 3°K. These ranges provide good optimization for bulk mixing of said reaction medium in balance with mixing power input from compression power input to recycle gas and/or mechanical agitation means. [00253] Said axial thermal gradient is defined herein as the temperature difference between the mass weighted average temperature of the 10% of said reaction medium at highest elevation, herein called the maximum temperature of said reaction medium, compared to mass weighted average temperature of the 30% of said reaction medium at lowest elevation, herein called the minimum temperature of said reaction medium, each value being time averaged over at least about 1 minute. [00254] The inventive gas flow rates and relatively small H/D disclosed herein are particularly useful for limiting thermal gradients. However, thermal gradients may be exacerbated by inapt placement of mechanical agitation devices, by flow impediments such as internal heat exchange surfaces and mechanical baffling means, and by inapt placement of gas and liquid feeding flows. [00255] An open (substantially empty of internal mechanical devices) BCR having a relatively short H/D as disclosed herein, using the inventively large gas flow ratios and superficial gas velocities herein, and provided with feeding locations of gas and liquid flows distributed azimuthally in the bottom 10% of said reaction medium is particularly useful for achieving the disclosed small values of axial thermal gradient. [00256] However, it is also possible to obtain a preferred thermal uniformity when using internal heat exchange surfaces and/or mechanical baffling means and/or mechanical agitation means. Yield losses of converted olefin into byproducts [00257] It is preferred to use various embodiments of the inventions herein to provide a very good utilization of said reaction medium and the recycle gas compression power to optimize the desired hydroformylation chemistry at a large yield balance compared to undesired byproduct reactions. Mass of heavy liquid byproducts/mass of aldehyde can be < 3%; < 2%; < 1% and < 0.5%. Moles of alkane byproducts/moles of olefin converted can be < 3%; < 2%; < 1%; and < 0.5%. Mass of total byproducts/mass of olefin converted can be < 4%; < 3%; < 2%; and < 1%. [00258] Said byproduct amounts comprise all byproducts made in a first main reaction medium plus byproduct amounts made in all processes associated with cooling said reaction medium, with separating a crude aldehyde product from catalyst of said reaction medium, and with purifying fresh reactant feeds in operations contacting the crude liquid aldehyde product. These byproduct amounts do not include additional amounts that are produced during distillation or other refining operation often used to provide a more highly purified liquid aldehyde product, e.g., after a crude liquid aldehyde product is exported from an aldehyde synthesis process. [00259] The formation rate of heavy byproducts is related strongly to the temperature, mass and aldehyde concentration of the liquid phase of said reaction medium, whereas the formation rate of aldehyde product relates to additional, separately controlled factors comprising hydroformylation catalysis, pressure, and aeration of said reaction medium along with the gas phase compositions of reactant gases. The formation rate of alkanes is affected strongly by catalysis, partial pressures of olefin and H2, and temperature of the reaction medium. Operation with relatively large formation rates of aldehyde product per unit mass of liquid phase of reaction medium is one aspect of suppressing the yield loss ratio of adverse byproduct formation, comprising heavy byproducts, alkanes, and total byproducts. [00260] Other important aspects for limiting adverse byproduct formation comprise controlling the absolute values and gradients of temperature and gas composition in a reaction medium; limiting the concentration of aldehyde within the liquid phase of a reaction medium; and substantially avoiding use of a heated, liquid-stripping separation of aldehyde outside of a first main reaction medium and vessel. The inventively larger gas flow ratios and other embodiments disclosed herein are useful for limiting adverse byproduct formation in all these regards. Syngas scrubber simplification and improvement [00261] Optionally, provide heating of a syn-gas scrubber process and vessel by heating at least a portion of a gas supply flow, rather than by heating a portion of crude liquid aldehyde. [00262] This differs from prior process design which similarly satisfies a preferred energy balance by heating a portion of liquid phase from scrubber underflow and pumping this heated liquid upwards and back into the scrubber at an elevated position. This steam heating of concentrated aldehyde liquid has at least three adverse effects, all avoided by instead heating a gas supply flow to the scrubber. Advantages are: [00263] It increases capital, operating and maintenance costs by requiring at least 1 additional liquid pumping operation; [00264] It increases the formation of heavy byproducts by forming an additional portion of steam-heated reaction medium; and [00265] It detracts from the optimal countercurrent sizing and operation of the contactor by refluxing underflow liquid back upwards in the scrubber vessel. [00266] As used herein, a syngas scrubber process and vessel comprise a means for contacting at least a portion of crude aldehyde liquid with at least a portion of H2, CO and/or combined syngas. [00267] It is preferred to use at least a portion of crude aldehyde liquid condensate from a recycle gas cooling condenser. It is preferred to use at least a portion of fresh makeup H2, CO and/or combined syngas. [00268] It is preferred that said contacting is substantially countercurrent, often with liquid flowing downwards by gravity through a rising gas flow, thereby recovering at least a portion of valuable olefin back into gas fed into a hydroformylation reaction medium and vessel and also transferring at least of portion of at least one non-inert impurity from the gas flow into exiting flow of crude liquid aldehyde. [00269] To improve phase-to-phase contacting and to promote countercurrent flow, it is preferred that said contacting includes mechanical staging means comprising any and all forms of liquid and gas distributors, tray- type staging means, packing-type staging means, and/or static-mixer-type staging means. [00270] It is preferred that at least a portion of total olefin dissolved in a liquid phase entering said scrubber is transferred into a gas phase exiting said scrubber and then fed into a hydroformylation reaction medium and vessel, thereby recovering valuable olefin out of a crude aldehyde condensate. Furthermore, it is preferred that the olefin concentration in at least a portion of crude aldehyde liquid exiting said scrubber is reduced as follows: 1) Moles of dissolved olefin per mole of crude liquid aldehyde ex scrubber < 80 ppm molar; 2) Moles of dissolved olefin per mole of crude liquid aldehyde ex scrubber < 40 ppm molar; 3) Moles of dissolved olefin per mole of crude liquid aldehyde ex scrubber < 20 ppm molar; and 4) Moles of dissolved olefin per mole of crude liquid aldehyde ex scrubber < 10 ppm molar. [00271] Besides recovering valuable olefin, it is preferred that at least a portion of at least one non-inert impurity present in a gas phase flow entering said scrubber is transferred either directly or in combination with chemical reaction into a liquid phase exiting said scrubber. [00272] Non-inert impurities are sometimes found in fresh makeup supplies of H2, CO, and/or syngas, and these can have severe financial consequences with respect to loss catalysis activity, even when present in very small concentrations. Their presence can even result in precipitation of catalyst metal from a reaction medium. Non-limiting examples of said non-inert impurities sometimes found in said makeup gas supplies comprise the following: O2, Amines, Certain sulfur compounds, e.g., hydrogen sulfide, carbonyl sulfide, Conjugated Dienes, e.g., 1,3-butadiene, Compounds with carbon-carbon triple bonds, e.g., acetylene, propyne, Phosphine gas, i.e., PH_3, certain transition metal-containing gases, e.g., AsH₃, nickel tetracarbonyl, iron pentacarbonyl, mercury. [00273] It is preferred that the temperature of a liquid phase exiting said scrubber process and vessel is within selected ranges. Preferred temperatures balance the equilibrium dissolution of olefin from liquid phase into gas phase and the reaction rates for non-inert impurities with aldehyde to form compounds more readily absorbed into the liquid phase, on the one hand, against the formation rates of heavy byproducts and against heating and/or cooling energy input requirements, on the other hand. Temperature of liquid phase exiting scrubber can be > 0°C; > 10°C; > 20°C; > 30°C. The temperature of liquid phase exiting scrubber can be < 100°C; < 90°C; < 80°C; and 8) < 70°C. [00274] The transfer of olefin dissolved in crude aldehyde condensate back into a gas phase essentially comprises a latent heat of evaporation and tends to cool the temperature of the scrubber operation. To maintain operating temperature of crude aldehyde liquid exiting said scrubber, it is often necessary to provide heating energy input associated with said scrubber operation, particularly in consideration of temperature ranges preferred for forming crude liquid aldehyde condensate. [00275] To control the temperature of a liquid phase exiting said scrubber process and vessel, it is preferred to heat at least a portion of at least one fed gas phase as follows. These ranges are in respect of useful utility fluid temperatures, useful metallurgy for handling H₂, CO, and/or syngas, preferred scrubber operating temperatures disclosed herein, and typical mass and enthalpy balances for countercurrent contacting of typical mass flows of fresh makeup feed gases H2, CO, and/or syngas combined with typical mass flow, composition, and temperature of crude liquid aldehyde condensate. The temperature of gas phase entering scrubber can be > 30°C; > 50°C; > 70°C; > 90°C; < 320°C; < 280°C; < 240°C, and < 200°C. [00276] This process embodiment eliminates 2 pumps on crude aldehyde liquid (duty and standby), and it replaces one steam-heated exchanger possibly abusing concentrated aldehyde liquid with one steam heated exchanger operating on syngas. It also reduces scrubber column height and diameter compared to the counterproductive backwards circulation of bottoms liquid. Combination with Purge Gas Reaction Process and Apparatus [00277] One potential disadvantage of providing the inventive, greater flows of gas through a reaction medium and vessel, pro rata to production rate, is a reduced conversion per pass of valuable olefin. In many cases (but not all, e.g. BISBI), this can lead to increased concentrations of olefin in the purge gas sourced from the recycle gas loop. [00278] In other cases, notably when one or more of the reactant gases is a partial poison for catalysis as well as reactant, e.g., BISBI, larger recycle gas flows through a reaction medium and vessel may enable activity to be sustained despite smaller concentrations of olefin because of providing more optimized CO and/or H₂ profiles by reducing axial gradients thereof. [00279] For cases where larger gas flow rates perhaps lead undesirably to increasing olefin concentration in purge gas flows, it is preferred to couple the present invention with a prior invention disclosed in IR-86251, namely a purge gas scrubbing reaction medium, vessel, and process. [00280] At least a portion of olefin in purge gas is converted to aldehyde in a liquid phase reaction akin to operation of a first main reaction medium and vessel, excepting that the purge gas flows upwards through the purge gas scrubber reaction medium and vessel only once, specifically without substantial recycle of exiting overhead gas, and wherein substantially all of the aldehyde produced in a purge gas scrubber reaction medium and vessel is removed from said medium and vessel by gas stripping in the exiting gas phase. [00281] More preferably, the purge gas reaction medium and vessel provide a relatively large conversion per pass of valuable olefin; and this is often achieved by using a relatively large ratio of H/D in a purge gas BCR, optionally using horizontal tray-staging means to minimize back-mixing of the gas phase downwards from nearer top toward nearer bottom, i.e. providing a gas phase Residence Time Distribution within said reaction medium and vessel that is a narrower RTD and more closely approaching plug flow of the gas phase. [00282] When at least one purge gas scrubbing reaction medium and vessel is used, it is preferred that the fraction of aldehyde produced in said scrubbing reaction medium and vessel compared to the total amount of aldehyde produced by one or more associated first main reaction mediums and vessels plus said scrubbing reaction medium and vessel lies within a range that is neither too small nor too large. [00283] When the production ratio from a purge gas scrubbing reaction medium and vessel is too small, it is likely not economical to provide said scrubbing reaction medium and vessel, even one as simple as a once-through- gas-flow BCR that is self-stripping of aldehyde product into exiting gas. [00284] When the production ratio from a purge gas scrubbing reaction medium and vessel is too large, the operation may cease to be self-stripping using once-through-gas-flow, unless augmented by inefficiently large amounts of additional, once-through H₂, CO and/or inert stripping gases. [00285] Accordingly, it is preferred the fraction of aldehyde produced in an associated second, purge gas scrubbing reaction medium and vessel is within selected ranges with respect to the amount of aldehyde produced in an associated first main reaction medium and vessel. Aldehyde moles from purge reaction/Aldehyde from main reaction can be > 0.1%; > 0.2%; > 0.3%; and > 0.4%. Aldehyde from purge reaction medium and vessel/Total Aldehyde can be < 18%; < 14%; < 10%; < 8%; < 4%; and <2%. [00286] Also, it is disclosed that one or more first main reaction mediums and vessels can provide purge gas for feeding into one or more second purge gas scrubber reaction mediums and vessels, and in all combinations thereof. In such cases, the above production ratios apply to the sum of all production from associated second purge gas reaction mediums and vessels to the sum of all aldehyde produced by the array of associated first main reaction mediums and vessels. [00287] Purge gas can be provided from any location within a recycle gas loop of a first main aldehyde production. [00288] More preferably, purge gas is provided downstream of a cooling condenser and after a preferred removal of cooled condensate, for this reduces the amount of aldehyde product sent with purge gas. [00289] It is still more preferred to provide purge gas from a position after the discharge from a recycle gas compressor, even though this subtracts from the amount of recycle gas provided by a fixed size of recycle gas compressor into the associated first main reaction medium and vessel. [00290] The benefit of providing a purge gas from the discharge of a recycle gas compressor is greater operating pressure within said purge gas reaction medium and vessel. [00291] A hydroformylation reaction is often strongly dependent on system pressure, with other factors such as gas compositions, temperature and catalysis of a reaction medium held about constant. [00292] Often the rate response of hydroformylation approaches between 2^nd and 3^rd power on system pressure, and a small percentage increase in system power can have double or triple that much effect on reaction rate. [00293] The increased hydroformylation activity provided by greater system pressure is particularly useful when trying to consume at least one of the reactants, often the olefin, toward virtually total conversion while also seeking to limit the mass of reaction medium and the physical size of the reaction vessel. [00294] Purge gas composition comprises olefin, CO, H2, and at least one inert impurity. Said inert impurity arises from fresh makeup reactant supplies and/or production from at least one chemical reaction within an aldehyde synthesis process. Typical inerts comprise alkanes, N2, CO2, H2O, alcohols, aldehyde condensation products, and amines. Volatile inerts tend to accumulate in the recycle gas composition when their solubility in an exiting crude liquid aldehyde product flow is smaller than the feeding plus production rate of said impurity; and a purge gas flow is required to limit the concentration of said volatile inert in recycle gas. [00295] This reduction in propylene content of the gas phase reduces the amount of propylene + CO + H2 purged per unit of inert gases purged, and it may also lead to a reduced hydrogenation yield loss of alkene to alkane. [00296] In addition, a cooler reaction temperature means less propylene pressure is required to stabilize BISBI. [00297] In combination, the need for less propylene concentration in a reaction medium and vessel leads to less purge gas value loss by operation at lower mole% propylene in the moles of gas purged, and it also allows operating at a higher mole% of propane in the purge gas thereby requiring fewer moles of purge gas flow. (Restating, the ratio of propylene/propane in purge gas is leveraged by reducing propylene concentration, enabled by reducing the axial gradient of CO, by cooling said reaction medium, and by increasing propane concentration in recycle gas; and the loss of CO/H2 is also reduced by increasing propane concentration in purge gas.) [00298] A cooler reaction medium and vessel leads to less hydrogenation rate of alkene to alkane directly leading to lower purge gas flow rate and less loss of value of alkene + CO + H₂ and/or less cost of their recovery. In addition, it is likely that the reduced concentration of propylene, which is enabled by a cooler BISBI reaction medium provided with reduce axial gradient of gas phase concentration of CO, is further helpful in reducing hydrogenation rate and purge yield losses and/or olefin recovery costs. Reduced capital costs: [00299] Speculatively, capital cost can be reduced by about 20% compared to a single standard (R-1 CSTR) reactor unit of comparable capacity. This is by elimination of equipment items in a revised, simplified process design and by using a simpler reaction vessel and is, furthermore, despite using greater flow rates of recycled gas. [00300] Speculatively, capital cost can be reduced by greater than about 35% when replacing 2 or more reactor units having conventional capacities of 150 to 225 KTA with a single larger reactor unit; and this can provide an essentially unlimited design capacity by using the simplified process design and despite using greater flow ratios of recycled gas. This is due to the economies of scale plus the revised, simplified process design and reaction vessel. [00301] Capital cost can also be reduced markedly on a scaled down or scaled up unit by fitting a reaction medium and vessel and other equipment sizes more aptly to production rate set by market demand and/or feed gas supply limits, as contrasted to installing an integer number of conventional Oxo R-1 type reactor units. [00302] The R-1 vessels cannot be scaled significantly with confidence (either up or down) due to profoundly important, but imperfectly known and understood, internal gradients in all 3 spatial dimensions for temperature, gas and liquid phase compositions, and gas-to-liquid mass transfer rates. Reduced utility costs: [00303] Despite the increased gas circulation rates needed in the following disclosures, a net reduction in utility costs is obtained as follows: Compared to present R-1 and gas recycle loop, this invention provides a more thermo- mechanically efficient compressor than presently used. It eliminates unnecessary gas loop dP by eliminating the gas-to-gas interchanger, which comprises 2 flow passes of recycle gas, eliminating the recycle gas preheater on the compressor suction by using a larger diameter and mechanically robust centrifugal impeller operating directly at motor speed, replacing the isenthalpic capacity throttling valve on the compressor discharge with isentropic guide vanes on the compressor suction, replacing, in the BCR option, the separate gas-liquid de-entrainment vessel for gas exiting overhead from a reaction vessel with a demister pad placed inside and near the top of the larger diameter reaction vessel. [00304] Embodiments of this invention eliminate the need for the capital and energy costs of an enclosed cooling water system, which is needed for cooling exchanger tubes located inside a reaction vessel in order to limit corrosion and fouling woes inside a reaction vessel. This applies for an option providing cooling of liquid phase of reaction medium outside a reaction vessel and/or for an option providing such a large gas flow rate that heat exchange surface in contact with liquid phase of reaction medium is obviated. [00305] Embodiments of this invention eliminate, in the BCR option, the utilities associated with reaction vessel agitator, i.e. main motor, gear-box lube oil utilities, shaft seal utilities. [00306] Embodiments of this invention avoid the incremental heating (steam), cooling (CWS), pumping and compression (electricity) costs that are required when using a liquid stripping system for aldehyde separation from reaction medium. [00307] This is particularly true for HBu because the relatively large solubilities of propylene and propane force the capital cost and power requirements for the multi-stage flash gas compressor associated with a liquid stripping unit to exceed the incremental cost and power requirements of a larger gas recirculation compressor on a main reaction unit. [00308] Steam heating of ligand and aldehyde containing liquids is eliminated, except for start-up, by avoiding a liquid stripping unit. [00309] Cooling water duties are minimized because liquid pump and their shaft power input are reduced compared to liquid stripping, because there is no steady state steam input for liquid stripping, because compression energy is reduced compared to liquid stripping, and because the mechanical agitator can be removed from a reaction vessel (BCR option). Minimized working capital for Rh inventory [00310] Liquid mass of reaction medium can be reduced in its ratio to aldehyde production rate compared to the existing R-1 type reactors which have unfortunately and undesirably large gradients in aeration, temperature and gas and liquid phase compositions of the 3 reactant gases. In addition to axial composition and thermal gradients discussed above, the CO and H₂ are so sparingly soluble that high energy input near impellers and in axial core of reaction medium cannot dissolve and store enough CO and H2 to support fully the intrinsic reaction rates radially outside and axially below the cooling tube bundles. [00311] Thus, the need for extra Rh mass (catalyst sites) and mass of liquid phase (gas dissolution sizing) can be reduced within the reaction medium and vessel of a gas-only stripped design; and this is in addition to avoiding liquid phase of reaction medium in a liquid stripping design. [00312] In addition, the mass of Rh and the mass of reaction medium can be sized more appropriately for the desired production capacity, compared to using than using an integer number of R-1 clone vessels. [00313] Accordingly, working capital is reduced, and this is in addition to reducing operating cost when mother liquor must be withdrawn periodically due to chemistry reasons (ligand degradation, heavies accumulation). [00314] This invention provides an improved process and apparatus for recovering valuable olefin, carbon monoxide (CO) and molecular hydrogen (H2) compounds from a purge gas stream using a metal-catalyzed (e.g., rhodium Rh) hydroformylation reaction process. More preferably, the purge gas stream feed is taken at least in part from a separate, though associated, metal-catalyzed (e.g., rhodium Rh) hydroformylation reaction process. The invention uses reactive scrubbing of olefin, CO and H2 in a separate hydroformylation reactor, reaction medium and process, more preferably a Bubble Column Reactor (BCR) vessel and process with inventive advantages. [00315] One aspect of the invention provides for removal of the preponderance, or even entirety, of the aldehyde produced in the purge gas scrubbing reactor and reaction medium by gas-phase stripping from the reaction medium, followed by cooling of the gaseous effluent to condense the valuable aldehyde, specifically limiting or even eliminating the following. The amount of effluent gas that is recirculated to the reactor and reaction medium, which is sometimes done to promote additional gas-stripping of aldehyde product, is greatly limited or even eliminated altogether. And, the amount of aldehyde recovered by liquid-phase stripping of aldehyde from the reaction medium is also greatly limited or even eliminated altogether. [00316] Liquid-phase stripping comprises removing a portion of the reaction medium from the reaction vessel and then processing this portion in an aldehyde recovery step comprising a) pressure reduction and/or heating of the removed liquid phase in order to flash additional aldehyde product, b) recovery of at least a portion of this flashed aldehyde by cooling and to produce a liquid phase condensate, and c) re-pressurizing at least a portion of the un-flashed liquid phase, which now contains a reduced fraction of produced aldehyde, and returning this back into a hydroformylation reactor and reaction medium. [00317] One aspect of the improvement is to provide a BCR that is designed to provide an improved approach toward plug-flow of the gas phase of the reaction medium even when the liquid phase of the reaction medium has a long residence time and is relatively greatly back-mixed. [00318] One aspect of the invention is to recover a very great fraction of olefin content of purge gas while limiting the production of an increased fraction of heavy condensation products of aldehyde that are themselves a yield loss and/or product purification burden. [00319] One aspect of the invention is using significantly different catalytic formulation in the purge gas scrubber BCR than in the main hydroformylation reactor, e.g. a markedly different liquid phase concentration of catalyst metal, e.g., Rh, and/or different ligand co-catalyst moiety. [00320] One aspect of the invention is to provide treatment in the purge gas scrubbing reactor of at least a portion of the mother liquor being permanently purged from a first hydroformylation reactor, the treatment providing additional recovery of aldehyde product from the first reactor and thereby increasing the concentration of catalyst metal prior to processing for metal recovery. [00321] Metal-catalyzed (e.g., rhodium, Rh) hydroformylation of olefins with CO and H2 is used commercially form aldehydes, e.g., hydroformylation of ethylene and propylene form propionaldehyde and butyraldehyde, respectively. Owing to the gaseous tendency of CO and H2, the reaction medium is typically at least a 2-phase reaction medium comprising a gaseous phase and at least one liquid phase. Furthermore, the reaction medium may sometimes comprise multiple liquid phases or even a solid phase. [00322] Reactive consumption of fed olefin, CO, and H2 leads to concentrating undesirable inert gaseous compounds in the gaseous phase of the reaction medium. Typical inert gaseous compounds comprise carbon dioxide (CO2), water, methane, ethane, propane, argon, and molecular nitrogen (N2), though many others also may be present. [00323] Some quantities of these undesirable inert gases are fed as relatively small concentrations of impurities within the commercial supplies of olefin, CO, H2 and synthesis gas (syngas). [00324] Additional quantities of some of these inert gases are formed in the reaction medium by unwanted side reactions, e.g. hydrogenation of olefins to alkanes. [00325] The kinetic needs of the gas-to-liquid mass transfer and of the liquid- phase hydroformylation reaction are such that the conversion per pass of at least one compound among the fed olefin, CO and H2 syngas is often much less complete than 100% per pass though the two-phase reaction medium. Accordingly, valuable feed compounds are often recovered as a gaseous mixture separated from at least one the various effluents from the hydroformylation reactor and reaction medium; and then these separated feed compounds, often in a mixture comprising all 3 hydroformylation reaction compounds, are recirculated (recycled) back to the reactor and reaction medium to supplement fresh feeds of olefin, CO, and H2. Such a recirculated gaseous mixture is often called “recycle gas”. [00326] Such recycle gas streams may be formed directly from a gaseous flow leaving the reactor vessel and multi-phase hydroformylation reaction medium by gravity separation, with gas mostly leaving overhead and liquid mostly remaining below; or it also may be formed by withdrawing a portion of the multi-phase hydroformylation reaction medium from the reaction vessel into a lower pressure flash separation step, which often comprises a heating means to add enthalpy to adjust the temperature in the flashing multi-phase mixture. In both cases, a recompression step is needed to increase the pressure of the respective recycle gas stream to a pressure sufficient to be reintroduced into the source reactor vessel and multi-phase hydroformylation reaction medium. [00327] Over time, the various undesirable inert gaseous compounds accumulate in various recycle gas flows of a hydroformylation reaction system, and some means must be provided for the removal of these inert gaseous compounds from the hydroformylation reaction process. [00328] A typical means of removing the accumulating inert gaseous compounds is via an intentional flow of a gaseous mixture divided from one or more of the recycle gas streams before a remaining portion of the recycle gas is returned to the source reactor. Such a gaseous flow divided from recycle gas being returned to the first reactor is often referred to as reactor purge gas or more simply purge gas leaving the primary hydroformylation reaction system. The flow rate of this purge gas relative to the feeding rate and formation rate of the inert gaseous compounds is used to adjust and control their steady state concentrations in the gaseous phase of the multi-phase hydroformylation reaction medium within the primary reactor. [00329] On the one hand, it is desirable to set the purge gas flow as large as possible to maintain relatively high concentrations of olefins, CO and H2 in the gas phase of the hydroformylation reaction medium. This improves the kinetics for gas-to-liquid mass transfer and for the liquid-phase reaction chemistry forming aldehydes. [00330] On the other hand, it is desirable to reduce the purge gas flow as low as possible because of the incumbent loss of valuable olefins, CO, and H2, wherein olefins often comprise greatest value. [00331] Thus, a purge gas flow rate is selected, adjusted and controlled to balance these two competing objectives of maintaining productivity and selectivity in the hydroformylation reactor and reaction medium while avoiding too much loss of valuable feed compounds. [00332] Physical separation of the purge gas completely into its molecular constituents in order to recover olefins, CO, and H2, although possible, is typically not commercially practicable. [00333] Since olefin is often the most valuable component of purge gas, considering both fed compound price and fraction remaining in purge gas, it is known in the art to recover a significant portion of valuable olefin from purge gas by physical separation means and by chemical reaction means. [00334] In particular, it is known in the art to recover a portion of the valuable olefin and also a portion of the syngas from purge gas by feeding purge gas to a second, separate hydroformylation reactor and reaction medium. Such a second reactor and second reaction medium converts into additional aldehyde products at least a portion of olefin, CO, and H2 that have already passed through a first reactor and reaction medium. (The terms second and first reactors and reaction media refer to the relative sequence of the two and do not preclude that there may be additional upstream and downstream hydroformylation reactors and reaction media.) [00335] This second reactor typically produces a much smaller quantity of aldehyde. Accordingly, the terms “scrubbing” and “polishing” are sometimes used to describe in lieu of “second” reactor and “second” reaction medium. [00336] Such a purge gas scrubbing reactor may have quite different mechanical size or construction compared to a first reactor. The various process steps associated with a scrubbing reactor may also be different or the same and may be closely integrated with or separated from process steps associated with a first reactor. [00337] Deficiencies in prior art for purge gas reactive scrubbing processes and reactors comprise the following, in various combinations: Some require withdrawal of a liquid phase from the purge gas scrubbing reactor and reaction medium substantially continuously in order to remove a significant portion of the produced aldehyde; and others require recycling of partially purified overhead gas from the scrubbing reactor back to the scrubbing reactor in order to remove a significant portion of the produced aldehyde. Some require relatively intimate comingling of the catalyst bearing mother liquor of the main hydroformylation reactor with the reaction medium in the purge gas scrubbing reactor. Some lack the ability to process mother liquor purged from the main reactor systems to obtain a more concentrated mother liquor, in order to recover additional aldehyde product and increase the concentration of catalyst metal before subsequent metal recovery processing. Some lack sufficient staging of the gas phase passing through the reaction medium toward a plug-flow residence time distribution. Some provide insufficient minimum residence time of the gas phase passing through the reaction medium. Some provide insufficient conversion per pass of olefin compound. Some produce too much heavy condensation products of aldehyde relative to aldehyde., Some are substantially more expensive to build and to operate than inventions herein. [00338] Referring to Figure 1, it shows a hydroformylation reactor process for converting at least a portion of purge gas feed stream 1 derived from at least one first hydroformylation reactor process into liquid crude aldehyde product 33 using a multi-phase hydroformylation reaction medium 11 that is at least partially contained within purge gas scrubbing reactor means 10. Optionally, reactor means 10 is referred to herein as a purge gas polishing reactor and a second reactor; and reaction medium 11 is referred to herein as a purge gas polishing reaction medium and a second reaction medium. It is preferred that the flows of streams 1 and 32, disclosed below, are substantially continuous in time preferably making reactor means 10 and reaction medium 11 continuous flow reactor means and reaction medium. [00339] The inventors contemplate that reactor means 10 may be one device, e.g. a single pressure vessel, length of conduit, and so on, essentially containing reaction medium; and the inventors also contemplate that reactor means 10 may be multiple such devices that are connected in series, in parallel, or in any combination wherein each reactor means essentially contains a portion of reaction medium 10. Thus, reaction medium 10 may be in one continuous mass or in multiple, physically separated portions. Preferred reactor types comprise all types of continuous flow reactors, with or without mechanical agitation, more preferred are continuous flow bubble column reactors, and most preferred are continuous flow bubble column reactors with gas-flow-staging internals (e.g. baffles, sieve trays, and so on). A preferred embodiment of reactor means 10 is described later in the figures. [00340] Stream 1 essentially comprises gaseous olefin, CO and H2 that have previously passed unreacted through at least one first hydroformylation reactor and reaction medium. In addition, stream 1 often comprises vaporized portions of CO2, N2, argon, alkanes, aldehydes, alcohols, amines, H2O, aldehyde condensation byproducts, and catalyst decomposition byproducts. Certain preferred ranges of the chemical composition of stream 1 are disclosed later herein. [00341] Stream 1 may be provided by a single flow conduit or by multiple flow conduits. The source of stream 1 may be a single first hydroformylation reactor and reaction medium; or stream 1 may derive from multiple first hydroformylation reactors and reaction media, even those operating at much different conditions from each other. When stream 1 derives from more than one first reactor and reaction medium, the constituent mass flows may be conducted separately into purge scrubbing reactor means 10 and reaction medium 11, or they may all be combined before entering reactor means 10 and reaction media 11, or they may be partly separated and partly combined; e.g., when 4 separate first reactors and reaction media each provide a portion of purge gas stream 1, all 4 constituents of stream 1 may be combined in a single conduit to form stream 1, or all 4 constituents of stream 1 may be separately conducted into reactor means 10 and reaction media 11, or 2 may be combined with each other and 2 may be left separated, and so on. [00342] Additional stream 2 is optional and may not be present. When present, it is preferred that the flow of stream 2 is substantially continuous in time. Stream 2 may be formed from any number of independent sources of gases and liquids in order to provide various composition ranges and flow ranges disclosed herein. Component compounds disclosed as particularly useful for stream 2 comprise alkanes, olefins, CO, H2, N2, and preferred compositions are disclosed later herein. Notably, stream 2 does not comprise significant amounts of hydroformylation catalyst compounds, which are provided in stream 3. The various constituent flows of stream 2 may be conducted singly or in any combination with each other before arriving at reactor means 10 and reactor medium 11. Furthermore, various constituent flows of stream 1 and stream 2 may be mixed with each other in any combination, or they may be left entirely separate before entering reactor means 10 and reaction medium 11. [00343] In a preferred embodiment of the invention, all constituents of stream 1 and of stream 2 are combined with each other in a single conduit before entering reactor 1 and reaction medium 11. In another preferred embodiment, all gaseous constituents of stream 2 are combined with stream 1 in single conduit, but liquid constituents of stream 2, e.g. liquid propylene, are conducted separately into reactor means 10 and reaction medium 11. [00344] In Figure 1, an optional cooling means 20 is provided, along with coolant supply 21 and coolant return 22, owing to the exothermic nature of the hydroformylation reaction. The exchange surfaces separating reaction medium 11 and coolant 21 and 22 within means 20 may be formed in any shape and may be constructed of any material suitable to the pressure, temperature, composition, and other considerations of the operating environment, as is known in the art. As a preferred embodiment, cooling means 20 comprises substantially planar and/or cylindrical surfaces separating reaction medium 11 and coolant 21 and 22, e.g. plate exchangers and tubular exchangers, and the surfaces are made of conductive and corrosion-resistant metals, especially grades of stainless steel, as is known in the art. [00345] The inventors contemplate that cooling means 20 may be a single device or multiple devices and that there may be various different sources of coolant, e.g. cooling water, brine, glycol solutions, ambient air, and so on, supplied independently to the various cooling means. The inventors further contemplate that cooling means 20 may be disposed directly inside reactor means 10 and reaction medium 11, as shown in the figures, or that cooling means 20 may located outside of reactor means 10 and outside or reaction medium 11 as shown in the figures. [00346] In FIG.1, a gas-liquid disengaging zone 14 is provided within reactor means 10 so that a spent gas stream 12 is withdrawn and conducted to a cooling, condensing and separating means 31 wherein cooling is sufficient to condense at least a portion of a liquid crude aldehyde product 33. [00347] Optionally, gas-liquid disengaging zone14 may not be present within reactor means When this is the case, a portion of reaction medium 11 from reactor means 10 is conveyed into a separate container for disengaging a spent gas stream in a gas liquid disengaging zone. [00348] Cooling and condensing means 31 is provided with coolant supply 34 and coolant return 35. The exchange surfaces separating spent gas 12 and coolant 34 and 35 within means 31 may be formed in any shape and may be constructed of any material suitable to the pressure, temperature, composition, and other considerations of the operating environment, as is known in the art. As a preferred embodiment, cooling means 31 comprises substantially planar and/or cylindrical surfaces separating gas and liquid stream 21, 32, and 33 from coolant 34 and 35, e.g. plate exchangers and tubular exchangers, and the surfaces are made of conductive and corrosion-resistant metals, especially grades of stainless steel, as is known in the art. [00349] Furthermore, the inventors contemplate that means 31 may comprise multiple heat exchange means operating with the same or even different cooling media, e.g. a first heat exchange means operating with cooling tower water coolant and a second heat exchange means operating with chilled water, brine, or glycol solution coolant, and so on. [00350] The inventors further contemplate that the means for separating a portion of liquid condensate, i.e., a portion of liquid crude aldehyde product stream 33, from the gaseous stream may be physically separate from the heat transfer means, e.g., a separate liquid knock-out separator vessel may be provided that is not integrally constructed with along with heat exchange surfaces. The remaining uncondensed gases and vapors exiting from means 31 are then discharged as vent gas stream 32. [00351] In addition to gaseous compounds, streams 1, 12, 30, and 32 may comprise small fractions of condensed liquids entrained as droplets, mists, and aerosols. For reasons of process efficiency and mechanical integrity, it is preferred that the mass fraction of liquids in streams 1, 30 and 32, each considered separately, be less than about 4%, more preferred less than about 2%, still more preferred less than about 1%, and most preferred less than 0.5%. [00352] Also in FIG.1, at least one source of liquid comprising catalyst compounds is provided via stream 3 entering reactor means 10 and reaction medium 11. According to one aspect of the invention, the time-averaged mass flow of stream 3 is essentially zero for a continuous 7-day period. According to another aspect of the invention, the flow of stream 3 is merely quite small in comparison to the flow of crude aldehyde product 33, notwithstanding start-up, shut-down and upset operating events. According to this latter aspect of the invention, it is preferred that the time-averaged mass flow of stream 3 is less than about 10%, more preferred less than about 8%, still more preferred less than about 6%, and most preferred less than 4% compared to the time- averaged mass flow of stream 33 during a continuous 7-day period. [00353] In a preferred embodiment of the invention, the flow of stream 3 is intermittent. According to this aspect of the invention, it is preferred that the flow of stream 3 is non-zero less than about 12%, more preferred less than about 8%, still more preferred less than about 4%, and most preferred less than 2% of the hours during a continuous 7-day period. According to this aspect of the invention, it is preferred that the number of distinct events wherein the flow of stream 3 is cycled on and off is less than about 12 per day, more preferred less than about 8 per day, still more preferred less than about 6 per day, and most preferred less than 4 per day. [00354] Similarly in FIG.1, at least one withdrawal flow of liquid comprising catalyst compounds is provided via stream 4 leaving reaction medium 11 and purge gas scrubbing reactor means 10. According to one aspect of the invention, the time-averaged mass flow of stream 4 is essentially zero for a continuous 7-day period. According to another aspect of the invention, the flow of stream 4 is merely quite small compared to the flow of crude aldehyde product 33, notwithstanding start-up, shut-down and upset operating events. According to this aspect of the invention, it is preferred that the time-averaged mass flow of stream 4 is less than about 5%, more preferred less than about 4%, still more preferred less than about 3%, and most preferred less than 2% compared to the time-averaged mass flow of stream 33 during a continuous 7- day period. [00355] In a preferred embodiment of the invention, it is preferred that the flow of stream 4 is intermittent. According to this aspect of the invention, it is preferred that the flow of stream 4 is non-zero less than about 12%, more preferred less than about 8%, still more preferred less than about 4%, and most preferred less than 2% of the hours in a continuous 7-day period. According to this aspect of the invention, it is preferred that the number of distinct events wherein the flow of stream 4 is cycled on and off is less than about 12 per day, more preferred less than about 8 per day, still more preferred less than about 6 per day, and most preferred less than 4 per day. [00356] In a preferred embodiment of the invention, the flow volumes and frequencies of streams 3 and 4 are adjusted such that the concentration of catalyst metal in supply stream 3, when it is flowing, divided by the concentration of catalyst metal in removal stream 4, when it is flowing, is kept within specified ranges. [00357] When stream 3 is relatively rich in aldehyde content, i.e., has a greater aldehyde fraction than found within reaction medium 11, it is preferred that the metal concentration ratio is at least about 1.1, more preferred at least about 1.2, still more preferred at least about 1.5, and most preferred at least 2. This situation may arise when the source of catalyst liquid stream 3 is from another hydroformylation reactor and comprises significant quantities of volatile aldehyde. In such cases, it is possible to adjust the flows and compositions of streams 1 and 2 such that spent gas stream 12 removes more aldehyde that is being produced within reaction medium 11. As a result, the purge gas scrubbing process is providing an additional function, namely concentrating a catalyst metal compound prior to a catalyst metal recovery step performed subsequently on stream 4 in preference to stream 3. In this embodiment, the integrated flow of exiting stream 4 will be corresponding less than the integrated flow of entering stream 3 when considering time periods at least about 7 days. [00358] On the other hand, the substantially continuous operation of reaction medium 11 may produce heavy condensation products of aldehyde in reaction medium 11 at a faster rate than they are removed via spent gas stream 12. In such a case, the mass and operating level of reaction medium 11 will rise and/or the catalyst system will be diluted. To avoid too much increase in level and/or too much loss of olefin conversion in this case, it is preferred that the catalyst metal concentration ratio is at least about 0.3, more preferred at least about 0.4, still more preferred at least about 0.5, and most preferred at least 0.6. In this embodiment, the integrated flow of exiting stream 4 will be corresponding more than the integrated flow of entering stream 3 when considering time periods at least about 7 days. [00359] The inventors have surprisingly discovered that it is possible to design and operate such a second purge gas scrubbing reactor, reaction medium and reactor process as shown in the figures employing a substantially or entirely one-pass-gas-flow-self-stripping mode wherein a substantial fraction, or even all, of the gas feeds pass through the reactor and reaction medium only once while still providing both conversion of a large fraction of the valuable olefin into aldehyde product and also removal from the reactor and reaction medium of a very large fraction, or even all, of the recovered aldehyde product as a vaporized component carried within the remaining unreacted gas phase. [00360] Thus, it is a key feature of one aspect of the invention herein that the preponderance, or even entirety, of aldehyde removed from reactor means 10 and reaction medium 11 is via spent gas stream 12 and liquid crude aldehyde stream 33 specifically without providing a liquid-phase removal and flashing system, as has been known in prior art. Thus, it is preferred that the amount of aldehyde in spent gas stream 12 is at least about 80%, more preferred at least about 90%, still more preferred at least about 95%, and most preferred at least 98% of all aldehyde removed from reactor means 10 and reaction medium 11. Also, it is preferred that aldehyde in liquid crude aldehyde stream 33 is at least about 75%, more preferred at least about 85%, still more preferred at least about 95%, and most preferred at least 95% of all aldehyde in the arithmetic sum of streams 33 and 4 during a continuous 7-day period. [00361] Furthermore, it is preferred that aldehyde in liquid crude aldehyde stream 33 is at least about 100%, more preferred at least about 103%, still more preferred at least about 106%, and most preferred at least 109% of the aldehyde formed within reaction medium 11. This recovered delivery in stream 33 of greater than 100% of the aldehyde produced within reaction medium 11, even while operating substantially or entirely with a one-pass-gas-flow-self- stripping mode is possible due to small portions of aldehyde often flowing into the reaction medium 11 via purge gas feed stream 1 coupled with a reduction in moles of gas phase leaving in vent gas stream 32 compared to feed stream 1. [00362] As defined herein, a liquid-striping step comprises removing a portion of the liquid phase of reaction medium 11 from reactor means 10 and passing this liquid phase into a step wherein additional quantities of liquid aldehyde are vaporized and thereafter condensed to form a portion of liquid crude aldehyde product that is substantially free of catalyst compounds. Such a liquid-stripping step often comprises a significantly reduced operating pressure compared to the source reaction medium and/or additional heating from a heating medium and/or additional sparging with a gaseous phase deficient in aldehyde vapor, in various arrangements, in order to vaporize an additional amount of aldehyde for separation from catalyst containing liquid. Such liquid stripping steps require considerable addition capital and operating expenses compared to the invention herein. [00363] However, the substantial absence of a liquid-stripping step for the present invention does not preclude that a small portion of product aldehyde may be recovered from the small flow of stream 4, when it is present according to some aspects of the disclosures herein. That is, small amounts of additional aldehyde product may recovered from stream 4 by liquid stripping as part of a catalyst metals recovery and isolation process providing that the catalytic transition metals contained therein are not thereafter returned directly to a hydroformylation reactor, e.g., within less than about 7 days. [00364] Furthermore, it is a key feature of one aspect of the invention herein that liquid crude aldehyde product 33 is provided without compressing a significant portion of vent gas 32 and recycling the portion back into multiphase reaction medium 11 in order to further increase the gas-stripping rate to match the aldehyde production rate. In this regard, it is preferred that the fraction of vent gas stream 32 that is compressed and recycled back to reaction medium 11 is less than about 50%, more preferably less than about 30%, still more preferably less than about 10%, and most preferably about nil, as is shown in FIG.1. [00365] Although recycling vent gas 33 potentially increases the gas-stripping of aldehyde from reaction medium and/or lowers the mole fraction of aldehyde within the liquid phase of reaction medium 11, such gas recycling is undesirable for other reasons. Recycling a portion of vent gas 32 back to reactor means 10 and reaction medium 11 requires gas recompression. This requires capital equipment and consumes energy. In addition, recirculation of a portion of vent gas 32 is often undesirable because this necessarily “back-mixes” and dilutes olefin concentration within at least a portion of reaction medium 11. This dilution of olefin concentration is often adverse to the efficient consumption of olefin to form aldehyde. Though the reaction kinetics for hydroformylation depends upon the specific catalyst mixture employed, the observed reaction order for olefin is very often a positive exponent meaning that lower concentrations of olefin very often lead to slower formation rates of aldehyde. Thus, vent gas recirculation often leads to larger volumes for reactor and reaction medium in order to obtain a desirably great conversion of olefin to aldehyde. This increase in size undesirably increases capital costs, and it also often leads to increased production of undesirable byproducts relative to the amount of aldehyde production. [00366] It is preferred that olefin portion in the sum of purge gas feed stream 1 plus additional feed supply stream 2 comprises principally light alpha-olefins. It is more preferred that that olefin portion comprises principally ethylene and propylene. The smaller molecular masses and the greater vapor pressures of produced propionaldehyde and butyraldehyde make them more amenable to gas-stripping. [00367] Furthermore, it is preferred that the mole fractions of olefin in the arithmetic sum of purge gas feed stream 1 plus additional feed stream 2 fall within selected ranges as now disclosed. It is preferred that the total olefin content and the ethylene content of the sum is less than about 30 mole %, more preferably less than about 20 mole %, still more preferably less than about 15 mole %, and most preferably less than 12 mole %. It is preferred that the propylene content of the sum is less than about 20 mole %, more preferably less than about 15 mole %, still more preferably less than about 12 mole %, and most preferably less than 9 mole %. Greater mole fractions of olefin increase the production rate of aldehydes and make the combinations of temperature, pressure and liquid phase composition needed within reactor means 10 and reaction medium 11 relatively inefficient for operating using a substantially or entirely one-pass-gas-flow-self-stripping mode. That is, olefin concentrations outside of preferred ranges tend to require reactor pressures that are too low or reactor temperatures that are too high leading to inefficiencies comprising excessive equipment sizing, excessive degradation of the catalyst system catalysts system, and/or excessive production of unwanted byproducts. [00368] On the other hand, when olefin fraction of the sum of purge gas feed stream 1 plus additional feed stream 2 is too lean, the heat of reaction from producing aldehyde in reactor means 10 and reaction medium 11 may be too little to make the reaction self-sustaining in balance with the cooling provided by purge feed gas stream 1 and additional feed stream 2, plus heat losses to ambient despite good insulation practice. Accordingly, it is preferred that the total olefin content of the sum is at least about 1 mole %, more preferably at least about 2 mole %, still more preferably at least about 4 mole %, and most preferably at least 6 mole %. [00369] The inventors have surprisingly discovered that judicious choices for operating one first reactor, reaction medium and reactor process can provide relatively efficient operation of the first reactor process while also providing suitable olefin concentrations, as disclosed above, in purge gas stream 1 without the need for significant added stream 2. In a one embodiment of the invention, the mass flow rate of added feed stream 2 is quite small compared to the mass flow of stream 1. In this embodiment, it is preferred that the mass flow of stream 2 is less than about 20%, more preferred less than about 10%, still more preferred less than about 5%, and most preferred effectively nil compared to the mass flow of stream 1. [00370] This is not an obvious optimization of the first reactor process operating conditions considered alone, i.e., without also considering the opportunity to convert olefin in purge gas stream 1 into additional aldehyde in a second reactor and reaction medium operated without a significant recycle gas step and without a significant liquid stripping step, as disclosed herein. [00371] The stoichiometric molar ratios of olefin, CO and H2 consumed in making aldehyde product is about 1:1:1, though the precise usage ratios in practice are often slightly variant owing to unwanted side reactions, e.g. hydrogenation of olefins to alkanes using H2 at a molar ratio of 1:0.5 without any consumption of CO. However, the composition of purge gas feed stream 1 may be dramatically different from 1:1:1 (olefin:CO:H2) depending upon the catalysis and desired reactivity and selectivity ratios provided in a first hydroformylation reactor, reaction medium, and reactor process. Examples of considerations in the first reactor and reaction medium and process providing a portion of stream 1 comprise: much smaller concentrations of CO may be appropriate in order to avoid an inhibiting effect in catalysis, or much greater concentrations of CO may be useful to adjust the normal to iso ratio during butyraldehyde production; much larger concentrations of H2 may be appropriate to drive mass transfer of this sparingly soluble gas into the dissolved liquid phase H2 at sufficient rates or much smaller concentrations of H2 may be useful to limit alkane formation; much larger concentrations of olefin may be useful to improve stability of the catalyst complex against degradation or to limit excessive byproduct formation rates, or much lower concentrations of olefin may allow greater accumulation of inert gaseous compounds within recycle gas of the first reactor process; greater or lesser recycle gas rates may be selected in the first reactor process according to the size of compressors available, cost of compression energy, and tolerance of the first reactor design for greater or lesser aeration, which affects the degree of difference between first reactor feed and effluent gas compositions; the purity of commercial supplies of olefin, CO and H2 are variable and affect the accumulation of inerts in recycle gas and purge gas of the first reactor process; and so on. [00372] In other aspects of the invention, purge gas from two or more first reactors, reaction media, and reactor processes operated quite differently are combined to provide a purge gas stream 1, either mixed in a single conduit or provided in various combinations separately to reactor means 10 and reaction medium 11 as previously disclose, the composition of which falls within the disclosed preferred ranges of olefin composition without significant flows of added feed stream 2. [00373] When suitable conditions cannot be procured in some combination of purge gases from first reactors, reaction media and reactor processes, the inventors have discovered that added feed stream 2 can be provided to obtain selected objects of the invention, namely stable operation of the purge gas scrubbing reactor, reaction media and reactor process using a substantially or entirely one-pass-gas-flow-self-stripping mode as disclosed above. [00374] When the composition of purge gas feed stream 1 is too lean in olefin, it is preferred that stream 2 comprise mass flow rates of olefins to bring the combined composition of streams 1 and 2, calculated whether these stream are actually mixed in one conduit or conducted separately into reactor means 10 and reaction medium 11, into the preferred ranges disclosed above. Olefin content of stream 2 may be in gaseous state, a liquid state or in some combination thereof, i.e. in 2 or more separate portions of stream 2 or even in a single multiphase flow of olefin. Furthermore, the additional mass flow of olefin provided by stream 2 may be added directly to reactor means 10 and reaction medium 11, or the additional mass flow may be combined with purge gas stream 1 before entering reactor means 10 and reaction medium 11, or in some combination thereof. [00375] When the composition of purge gas feed stream 1 is too concentrated in olefin, it is preferred to add CO, H2, syngas, and/or inert gaseous compounds as components in stream 2 to bring the combined composition of streams 1 and 2, calculated whether these streams are actually mixed in one conduit or conducted separately into reactor means 10 and reaction medium 11, into the preferred ranges disclosed above. Furthermore, the additional mass flow provided by stream 2 may be added directly to reactor means 10 and reaction medium 11, or the additional mass flow may be combined with purge gas stream 1 before entering reactor means 10 and reaction medium 11, or in some combination thereof. [00376] CO, H2 and syngas are often sufficiently lower in value compared to olefin and aldehyde that economics are improved by intentionally placing greater mass flow rates of CO and H2 content in vent gas stream 32, even stream 32 is sent as a waste effluent stream to an environmental control apparatus, such as a flare or other oxidative destruction. [00377] Preferred inert gaseous compounds for diluting olefin composition into the disclosed ranges comprise N2 and methane. Methane is often relatively inexpensive compared to olefin, and methane provides useful energy content when vent gas stream 33 is being sent to a burner for heat generation, e.g. a steam boiler and possibly when vent gas stream 33 is being sent to an environmental control device for oxidative destruction, e.g., flaring, regenerative thermal oxidation, catalytic oxidation, etc., wherein control of the energy density of the flared gas must sometimes be controlled in selected ranges for safety or environmental reasons. [00378] In some cases, the molar mass flow of olefin in purge gas stream 1 may exceed the molar mass flow of CO, of H2 or of both in stream 1. To maximize conversion of olefin to aldehyde according to one aspect of the invention, it is preferred that stream 2 comprise mass flow rates of CO, H2, and/or syngas sufficient to bring the combined composition of streams 1 and 2 into preferred ranges, calculated whether these stream are actually mixed in one conduit or conducted separately into reactor means 10 and reaction medium 11. It is preferred that the combined composition of stream1 and stream 2 comprise a ratio of CO to olefin that is at least about 1.0, more preferably at least about 1.1, still more preferably at least about 1.5, and most preferably at least 2.0. It is preferred that the combined composition of stream1 and stream 2 comprise a ratio of H2 to olefin that is at least about 1.0, more preferably at least about 1.1, and still more preferably at least about 2, and most preferably at least 4.0. [00379] According to aspects of the invention herein, it is preferred that the conversion of olefin content in purge gas feed stream 1 is very great, leaving relatively little olefin in spent gas stream 12 and in vent gas stream 32. Accordingly, it is preferred that olefin concentration in stream 12 is less than about 8 mole %, more preferred less than about 4 mole %, still more preferred less than about 2 mole %, and most preferred less than 1 mole %. [00380] Owing to the removal of condensed crude aldehyde product stream 33, olefin concentration in stream 32 is often somewhat higher than in stream 30, notwithstanding that some olefin is also dissolved into the liquid phase of stream 33. Accordingly, it is preferred that olefin concentration in stream 32 is less than about 8.8 mole %, more preferred less than about 4.4 mole %, still more preferred less than about 2.2 % about, and most preferred less than 1.1 mole %. [00381] In order to achieve such small concentrations of olefin in streams 12 and 32, it is necessary that the reactive consumption of olefin is relatively large within reactor means 10 and reaction medium 11, notwithstanding operating substantially or entirely in one-pass-gas-flow-self-stripping aspects of the invention. Accordingly, it is preferred that at least about 70 mole %, more preferred at least about 85 mole %, still more preferred at least about 95 mole %, and most preferred at least 98 mole % of olefin in streams 1 and 2 (arithmetic sum) is reacted to form aldehyde or byproducts within reactor means 10 and reaction medium 11. [00382] It is preferred that most olefin is converted to aldehyde also is recovered as crude aldehyde product stream 33. In addition to reactive conversion of olefin into not-olefin, this also requires that byproduct formation is suppressed by suitable design of reactor means 10 and reaction medium 11, as disclosed herein, and that sufficient fractions of aldehyde are condensed by cooling and condensing means 30. [00383] Thus, it is preferred that the formation of alkanes from olefin within reactor means 10 and reaction medium 11 is largely suppressed. Using designs for reactor means 10 and reaction medium 11 as disclosed herein, it is preferred that the molar flow rate of any single alkane, i.e. ethane, propane, and so on, in streams 3 and 12 (arithmetic sum) is increased above the molar flow rate of the corresponding alkane in streams 1 and 2 (arithmetic sum) by less than about 4%, more preferably less than about 3%, still more preferably less than about 2%, and most preferably less than 1% times the molar flow rate of olefin in streams 1 and 2 (arithmetic sum). [00384] Thus, it is preferred that the formation of heavy condensation products of aldehydes within reactor means 10 and reaction medium 11 is largely suppressed. Using designs for reactor means 10 and reaction medium 11 as disclosed herein, it is preferred that the mass flow rate of heavy condensation products of aldehydes in streams 3 and 12 (arithmetic sum) is increased above the mass flow of the heavy condensation products of aldehydes in streams 1, 2, and 4 (arithmetic sum) by less than about 4%, more preferably less than about 3%, still more preferably less than about 2%, and most preferably less than 1% times the mass flow rate of olefin in streams 1 and 2 (arithmetic sum). [00385] Also, it is preferred that the loss rate of olefin to the combination of alkanes and heavy condensation products of aldehydes is largely suppressed. Using designs for reactor means 10 and reaction medium 11 as disclosed herein, it is preferred that the combined mass flow rate of ethane plus propane plus heavy condensation products of aldehydes in streams 3 and 12 (arithmetic sum) is increased above the mass flow rate of ethane plus propane plus heavy condensation products of aldehydes in streams 1, 2, and 4 (arithmetic sum) by less than about 4%, more preferably less than about 3%,still more preferably less than about 2%, and most preferably less than 1% times the mass flow rate of olefin in streams 1 and 2 (arithmetic sum). [00386] In order to improve the recovery efficiency of condensed aldehyde, it is preferred that the minimum pressure and minimum temperature of the aldehyde containing flows obtained within cooling and condensing means 31 are within selected ranges, since greater pressures and cooler temperatures tend to promote condensation of more liquid crude aldehyde product stream 33. Thus, it is preferred that the minimum pressure of the gaseous process flow within means 31 is at least about 300 kPa, more preferred at least about 600 kPa, still more preferred at least about 1,000 kPa and most preferred at least 1,800 kPa. (44, 87, 145, 261 psi) It is preferred that the minimum temperature of the gaseous process flow within means 31 is less than about 40°C, more preferred less than about 30°C, still more preferred less than about 25°C, and most preferred less than 20°C. This minimum process temperature within means 31 does not preclude that the temperature of vent gas stream 32 exiting means 31 may be warmer due to the presence of an economizing interchanger within means 31. By suitable combinations of pressure and temperature it is preferred that at least about 80 mole%, more preferred at least about 85 mole %, still more preferred at least about 90 mole%, and most preferred at least 95 mole % of the aldehyde content within stream 12 is produced within liquid crude aldehyde stream 33, with the balance of the aldehyde moles being found within vent gas stream 32. [00387] Because purge feed gas stream 1 often contains vaporized aldehyde and because the molar flow rate of vent gas stream 31 may often be reduced significantly below the molar flow rate of stream 1, it is often possible to remove into liquid crude aldehyde stream 33 a portion of the aldehydes found within purge feed gas stream 1. Accordingly, it is preferred that the moles of aldehyde recovered in stream 33 divided by the moles of olefin fed in streams 1 and 2 (arithmetic sum) is at least about 84%, more preferably at least about 88%, still more preferably at least about 92%, and most preferably at least 96%. [00388] To achieve the simultaneous objectives inherent with very conversion of olefin fed in streams 1 and 2 into aldehyde product condensed in stream 33 while operating substantially or entirely in an one-pass-gas-flow-self- stripping aspect of the invention, selected ranges of reactor pressure and temperature are preferred. [00389] In the case of reaction temperature, hotter temperature often tends to increase desirably the hydroformylation reaction rate and also to promote vaporization of aldehyde product. However, hotter temperature often tends to promote undesirably, even excessively, the decomposition rate of catalyst ligand compounds and the rate of formation of unwanted reaction byproducts, e.g. alkanes from hydrogenation and heavy condensation products of aldehyde, which may adversely affect operating economics in many ways. Thus, a balance must be found in reaction medium temperature between too cold and too hot. [00390] Accordingly it is preferred that the temperature within reactor means 10 and reaction medium 11 is at least about 60°C, more preferred at least about 70°C, still more preferred at least about 75°C, and most preferred at least 80°C; and it is also preferred that the temperature within reaction medium 11 within reactor means 10 is less than about 140°C, more preferred less than about 130°C, still more preferred less than about 125°C, and most preferred less than 120°C. [00391] Furthermore, the balancing of economic of objectives in purge gas scrubbing reactor means 10 and reaction medium 11 is often different than in at least one first hydroformylation reactor providing at least a portion of purge gas feed stream 1. According to this aspect of the invention, it is preferred that the maximum temperature within reactor means 10 and reaction medium 11 is hotter than the maximum temperature in the first reactor and reaction medium by at least about 2°K, more preferred at least about 4°K, still more preferred at least about 6°K, and most preferred at least 8°K. When reaction medium 11 comprises at least one catalyst compound also found in the first reaction medium, it is also preferred that the maximum temperature with reactor means 10 and reaction medium 11 is hotter than in the first reactor and reaction medium by less than about 40°K, more preferred less than about 30°K, still more preferred less than about 25°K, and most preferred less than 20°K. [00392] In the case of reaction pressure, greater pressure often tends to increase desirably the hydroformylation reaction rate, particularly since the mole fraction of olefin is preferably being greatly reduced to relatively small concentrations during passage through reactor means 10 and reaction medium 11; but greater reaction pressure also tends to suppress the desirable vaporization of aldehyde into stream 30. Accordingly, it is preferred that the pressure within reactor means 10 and reaction medium 11 and of stream 12 upon leaving reactor means 10 is at least about 300 kPa, more preferred at least about 600 kPa, still more preferred at least about 800 kPa, and most preferred at least 1,000 kPa (44, 87, 116, 145 psi); and it is also preferred that the pressure within reactor means 10 and reaction medium 11 and of stream 12 upon leaving reactor means 10 is less than about 6,000 kPa, more preferred less than about 4,000 kPa, still more preferred less than about 3,200 kPa, and most preferred less than 2,400 kPa (870, 580, 464, 348 psi). [00393] Furthermore, the balancing of economic of objectives in purge gas scrubbing reactor means 10 and reaction medium 11 is often different than in at least one first hydroformylation reactor providing at least a portion of purge gas feed stream 1. [00394] In one aspect of the invention, purge gas feed stream 1 is provided from at least one first hydroformylation reactor process providing at least a portion of stream 1 without compression by any means (e.g. mechanically driven by shaft work input, hydrodynamic such as educator or similar device, and so on) after leaving the first reactor and first reaction medium and before entering reactor means 10 and reaction medium 11. When this is the case, it is preferred that the minimum pressure within reactor means 10 and reaction medium 11 is less than the minimum pressure in the first reactor and reaction medium by at least about 40 kPa, more preferred at least about 60 kPa, still more preferred at least about 80 kPa, and most preferred at least 100 kPa (6, 9, 12, 15 psi); and it also preferred that the pressure difference is less than about 1,200 kPa, more preferred less than about 900 kPa, still more preferred less than about 600 kPa, and most preferred less than 300 kPa (174, 131, 87, 44 psi). [00395] In another aspect of the invention, at least a portion of a purge gas stream from at least one first reactor process providing at least a portion of stream 1 is compressed by some means after leaving the first reactor and first reaction medium and before entering purge gas reactor means 10 and reaction medium 11. When the compression occurs, it is preferred that the minimum pressure within reactor means 10 and reaction medium 11 minus the minimum pressure within the first reactor and reaction medium is at least about -90 kPa, more preferred at least about -60 kPa, still more preferred at least about -30 kPa, and most preferred at least 0 kPa (-13, -9, 0 psi). With the compression, it also preferred that the pressure difference is less than about 1,600 kPa, more preferred less than about 1,200 kPa, still more preferred less than about 800 kPa, and most preferred less than 400 kPa (232, 174, 116, 58 psi). [00396] Another object and aspect of the invention is that olefin content of vent stream 32 is quite small and is not particularly worth recovering as olefin chemical value. Accordingly, it is preferred that less than about 20 mole%, more preferred less than about 15 mole%, still more preferred less than about 8 mole %, and most preferred less than 4 mole% of olefin content of streams 1 and 2 (arithmetic sum) remains as olefin content in stream 12 or in stream 32, despite operating with a substantially or entirely one-pass-gas-flow-self- stripping mode for passage of the gaseous phase through reaction medium 11. In addition, it is preferred that at least about 50 mole %, more preferably at least about 30 mole%, still more preferably at least about 10 mole %, and most preferably effectively all of olefin content in stream 32 is not subsequently converted to aldehyde in a hydroformylation reactor. [00397] A related object and aspect of the invention is that vent stream 32 is sent either to an environmental processing step (e.g., flare, regenerative thermal oxidation unit, catalytic oxidation unit, and so on) or more preferably to an oxidative combustion device for recovery of thermal energy (e.g., a steam boiler, a heat-transfer medium heater, and so on) wherein the preponderance of the H2 and carbon containing compounds within stream 32 are converted to CO2 and H2O. [00398] Furthermore, it is preferred that the pressure of vent stream 32 is large enough not to need an additional compression step before oxidation, notwithstanding that there may be pressure and flow regulating devices, e.g. control valves, orifices, and so on, plus impurity removal devices, e.g. filters, scrubbers, demisters, and so on before vent gas stream 32 reaches the oxidation step, e.g. fuel burner, catalyst bed, and so on. Accordingly, it is preferred that the pressure of vent gas stream 33, measured immediately upon exit of means 31, i.e. before any back pressure and flow regulating devices, e.g. control valves, orifices, and so on, and before any additional impurity removal devices in the gaseous stream, e.g. filters, scrubbers, demisters, and so on, is at least about 300 kPa, more preferred at least about 600 kPa, still more preferred at least about 800 kPa, and most preferred at least 1,000 kPa. (44, 87, 116, 145 psi). [00399] Simultaneous attainment of conversion, yield and recovery objectives while employing a substantially or entirely one-pass-gas-flow-self- stripping mode of operation as disclosed herein for a purge gas scrubbing reactor requires certain design features that are perhaps never preferred for efficient production of larger quantities of aldehyde. That is, the design and operation of a second purge gas scrubbing reactor, reactor medium and reactor process are often inapt in at least one significant way for efficient design and operation for the bulk of commercial aldehyde production. For example, design features optimal for purge gas scrubbing may be greatly inapt for producing a desired ratio of normal versus branched aldehyde isomers for the primary production of butyraldehyde. For example, design features for an optimal purge gas scrubbing reactor and reaction medium are greatly inapt for the bulk of aldehyde production with regard to the size, geometry and/or mechanical construction of a first reactor with respect to capital cost, catalyst working capital, catalyst degradation losses, and byproduct formation rates. Accordingly, it is preferred that the amount of aldehyde obtained in liquid crude aldehyde product stream 33 is less than about 18 mole %, more preferred less than about 15 mole %, still more preferred less than about 12 mole %, and most preferred less than 9 mole % of aldehyde recovered in all first reactor processes (arithmetic sum) from which at least a portion of feed gas stream 1 is derived. [00400] On the other hand, apt provision of a purge gas scrubbing reactor, reaction medium and reactor process as disclosed herein relaxes constraints concerning olefin conversion rates in a first reactor, reaction medium, and reactor process, i.e. relaxes economic constraints on the economic consequences of purge gas sent from a first reactor, reaction medium and reactor process to an alternative disposition. Accordingly, it is preferred that the amount of aldehyde obtained in liquid crude aldehyde product stream 33 is at least about 1 mole %, more preferred at least about 2 mole %, still more preferred at least about 3 mole %, and most preferred at least 4 mole % of aldehyde recovered in all first reactor processes (arithmetic sum) from which at least a portion of feed gas stream 1 is derived. [00401] It is also preferred that the volume of reactor means 10 and reaction medium 11 are smaller compared to the sum of first reactors and first reaction medium (arithmetic sum) from which at least a portion of feed gas stream 1 is derived, notwithstanding that the reduced and declining concentrations of olefin in reactor means 10 and reactor 11 are often adverse to olefin reaction rate and olefin conversion rate therein, compared to first reactors and reaction medium. It is preferred that the volume of reactor means 10 is less than about 30%, more preferred less than about 25%, still more preferred less than about 20%, and most preferred less than 15% of the volume of all first reactors (arithmetic sum) from which at least a portion of feed gas stream 1 is derived. It is preferred that the volume of reaction medium 11 is less than about 25%, more preferred less than about 20%, still more preferred less than about 15%, and most preferred less than 10% of the volume of all first reactors (arithmetic sum) from which at least a portion of feed gas stream 1 is derived. [00402] The inventors have discovered that controlling “residence time distribution” (“RTD”) of the gas phase of reaction medium 11 into certain forms can improve the efficiency of purge gas reactor means 10 and reaction medium 11 to obtain great conversions of olefin despite small and declining concentrations of thereof. This is notwithstanding that similar gas-phase RTD forms as are preferred for purge gas reactor means 10 and reaction medium 11 are often greatly disadvantaged economically when employed in a first reactor and first reaction medium. For example, very large reactor sizes and very inefficient chemistry results when all hydroformylation is attempted in a well-mixed CSTR using very small concentrations of gaseous provided in spent gas stream 12 according to embodiments herein. [00403] In order to obtain such great conversions of olefin to aldehyde efficiently for a purge gas scrubbing application, it is preferred that the gas- phase of reaction medium 11 not be greatly back-mixed as would occur in a well-mixed stirred tank reactor, or even in a series of 2 well-mixed stirred tank reactors. Instead it is preferred that the residence time distribution (RTD) of the gas phase of reaction medium 11 more nearly approaches a plug-flow RTD. [00404] As known in the art and as defined herein, “CSTR” means a Continuous-flow, mechanically-Stirred Tank Reactor. As known in the art and as defined herein, a “well-mixed CSTR” means a CSTR operating with sufficient agitation such that the composition of each phase of reaction medium contained therein is effectively uniform throughout the reaction medium. [00405] The principles of RTD for mass flow through a chemical reactor and their utility in chemical reactor design and operation are well established. See, for example, Chemical Engineering Kinetics, J. M. Smith, second edition 1970, McGraw-Hill, especially chapter 6, "Deviations from Ideal Reactor Performance.” A residence time distribution (RTD) function is defined and described on pages 246 ff. therein. A well-mixed single tank reactor, often called a well-mixed continuous flow stirred tank reactor (CSTR) is one idealized case. Another idealized case for flow behavior is plug flow, sometimes called tubular flow or piston flow, where there is negligible convective mixing of mass with surrounding mass while flowing through a reaction zone. Methods for determining experimentally the residence time distribution function for actual, physical reaction zones are disclosed on pages 248 ff of Smith. The methods include introducing step inputs and/or pulse inputs of an inert tracer compound into the flow entering a reaction zone and then measuring the mass of the tracer exiting the reaction zone as a function of time. In recent years, using step and/or pulse inputs of a radioactive tracer material has proven particularly useful, in part because radioactive measurements on exiting flow provide a continuous, non-invasive determination of the mass of tracer exiting as a function of time. Acquisition of such data and reconstruction of the RTD function, including calculation of the mass-averaged residence time, using radioactive tracer methods are available on a commercial, contractual basis from multiple contractors, including for example Tracerco (Pasadena, Texas) and Quest TruTec (La Porte, Texas). [00406] In the following disclosure, a notation is adopted wherein "t" is time; the residence distribution function of time "J(t)" is the cumulative fraction of mass initially supplied to a phase of the reaction zone at time t=0 that then exits the reaction zone before time t; "tavg" is the mass-averaged residence time determined from J(t); "t/tavg" is reduced time meaning time divided by mass- averaged residence time; and "CMF(t/tavg)" is the residence distribution function of reduced time. For example, CMF(0.2) is the cumulative mass fraction initially supplied to a phase of the reaction zone at time t=0 that then exits the reaction zone before a reduced time of 0.2. The mass average residence time (tavg) of an aliquot of mass initially fed to an enclosure at time t=0 is calculated as [(t)*(mass of the aliquot exiting at time t)]/(total mass of the aliquot) integrated from time zero until at least about 98 percent of the mass of the aliquot has exited the enclosure. The units of tavg are simply any unit of time, and t/tavg is dimensionless. [00407] Thus, it is preferred that the CMF(0.2) of the gas phase flowing through reactor means 10, as measured by the flow of at least 1 inert gaseous compound within streams 1 and 2, that exits reactor means 10 is less than about the CMF(0.2) of 2 equally-sized well-mixed CSTRs arranged in series and having the same total reactor volume as reactor means 10, more preferred less than about CMF(0.2) of 4 equally-sized well-mixed CSTRs arranged in series and having the same total reactor volume as reactor means 10, still more preferred less than about CMF(0.2) of 6 equally-sized well-mixed CSTRs arranged in series and having the same total reactor volume as reactor means 10, and most preferred less than CMF(0.2) of 8 equally-sized well-mixed CSTRs arranged in series and having the same total reactor volume as reactor means 10. It is preferred that the reduced time t/tavg at which gas phase CMF(t/tavg) equals or exceeds 0.9 (90%) as measured by the flow of at least 1 inert gaseous compound within streams 1 and 2 flowing through reactor means 10 is less than about 1.95, more preferred less than about 1.68, still more preferred less than about 1.55, and most preferred less than 1.48. [00408] Thus, it is preferred that the CMF(0.2) of the gas phase flowing through reaction medium 11, as measured by the flow of at least 1 inert gaseous compound within streams 1 and 2, that exits reaction medium 11 is less than the CMF(0.2) of 2 equally-sized well-mixed CSTRs arranged in series and having the same total reactor volume as reactor means 10, more preferred less than the CMF(0.2) of 4 equally-sized well-mixed CSTRs arranged in series and having the same total reactor volume as reactor means 10, still more preferred less than the CMF(0.2) of 6 equally-sized well-mixed CSTRs arranged in series and having the same total reactor volume as reactor means 10, and most preferred less than the CMF(0.2) of 8 equally-sized well-mixed CSTRs arranged in series and having the same total reactor volume as reactor means 10. It is preferred that the reduced time t/tavg at which gas phase CMF(t/tavg) equals or exceeds 0.9 (90%) as measured by the flow of at least 1 inert gaseous compound within streams 1 and 2 passing through reaction medium 11 is less than about 1.95, more preferred less than about 1.68, still more preferred less than about 1.55, and most preferred less than 1.48. By means later herein, such preferred gas-phase distributions can be provided within a continuous reaction medium 11 contained within a single reactor means 10. [00409] Referring now to FIG. 2, a simplified process flow diagram shows another one of an unlimited number of possible optional arrangements within the ambit already disclosed for FIG.1, e.g., relating to possibly having multiple, possibly different reactors 10, reaction media 11, gas-liquid disengaging zones 14, cooling means 20, and condensing means 30 arranged in various parallel and series arrangements. [00410] In FIG.2, there are 3 physically separated and mechanically distinct units, i.e., Unit A, Unit B, and Unit C, assembled conforming to disclosure already made herein. All of the numbering shown in FIG. 2 has the same descriptions as employed for corresponding numerals of FIG.1, and the suffix letters in FIG.2 denote the Unit Identifiers A, B and C. For example, streams 1A, 1B and 1C are purge gas feed streams for Units A, B, and C, respectively, and conform to the disclosures for purge gas feed stream 1; reactors 10A, 10B, and 10C are reactors for Units A, B, and C, respectively, and conform to the disclosure for reactor means 10; and so on. Each of streams with corresponding numerals 1, 2, 3, 4, 12, 21, 22, 34 and 35 and shown separately for Units A, B, and C may have the same or different flows, temperatures, pressures and compositions according to the range of disclosures herein. Each of means with corresponding numerals 10, 20, and 30 may be physically similar or very different according to the range of disclosures herein. [00411] In FIG.2, observe that Unit A lacks an additional feed stream 2, an option already disclosed. Observe that Unit A lacks a gas-liquid disengaging zone and operates effectively completely filled with reaction medium 11A, an option already disclosed. In Unit A, a portion reaction medium 11A is conveyed effectively continuously into reactor Unit C by the unreacted gas phase portions of streams 1A and 2A. Observe that cooling means 20 and condensing means 30 may be provided singly or multiply for a particular Unit, as already disclosed, e.g., see Unit B and Unit C. Observe that cooling means 20 is sometimes disposed within reaction medium 11 and reactor means 10 and other times outside reaction medium 11 and reactor means 10, as already disclosed, e.g. see Unit B. Observe that reaction medium 11A in Unit A is mechanically agitated by agitator 16A in addition to gas-flow agitation, as already disclosed, but that Units B and C have only gas flow agitation. Furthermore, there are an unlimited number of other possible arrangements within the ambit of the present disclosure wherein a substantially or entirely one-pass-gas-flow-self-stripping mode is achieved. [00412] Referring now to FIG.3, a preferred embodiment of the invention is presented wherein reactor 20 is specifically a single bubble column reactor containing the preponderance of reaction medium 10. [00413] As defined herein, the term "bubble column reactor" shall denote a reactor for facilitating chemical reactions in a multi-phase reaction medium, wherein agitation of reaction medium 11 is provided primarily by the upward movement of gas bubbles through reaction medium 11. As defined herein, the term "agitation" shall denote work dissipated into reaction medium 11 causing fluid flow and/or mixing. As defined herein, the terms "majority", "primarily", and "predominately" shall mean more than 50 percent. As defined herein, the terms “preponderance” and “preponderantly” shall mean more than 70 percent. As defined herein, the term "mechanical agitation" shall denote agitation of reaction medium 11 caused by physical movement of a rigid or flexible element(s) against or within reaction medium 11. For example, mechanical agitation can be provided by rotation, oscillation, and/or vibration of internal stirrers, paddles, vibrators, or acoustical diaphragms located in reaction medium 11. As defined herein, the term "flow agitation" shall denote agitation of reaction medium 11 caused by high velocity injection and/or recirculation of one or more fluids in reaction medium 11. For example, flow agitation can be provided by nozzles, ejectors, and/or eductors. [00414] In a preferred embodiment of the present invention, less than about 50 percent of the agitation of reaction medium 11 in the bubble column reactor during oxidation is provided by mechanical and/or flow agitation, more preferably less than about 30 percent of the agitation is provided by mechanical and/or flow agitation, still more preferably less than about 10 percent of the agitation is provided by mechanical and/or flow agitation, and most preferably less than 5 percent of the agitation is provided by mechanical and/or flow agitation. It is preferred that the amount of mechanical and/or flow agitation imparted to the multiphase reaction medium is less than about 3 kilowatts per cubic meter of reaction medium 11, more preferred less than about 2 kilowatts per cubic meter, still more preferred less than about 1 kilowatts per cubic meter, and most preferred less than 0.5 kilowatt per cubic meter. [00415] Referring still to FIG. 3, a preferred bubble column reactor 20 is illustrated as comprising a vessel shell 200 having a reaction section 201 and a disengagement section 202. Reaction section 201 defines an internal reaction zone 208, while disengagement section 202 defines an internal disengagement zone 209. A predominately gas-phase feed stream 55 is fed through conduit 56 and is introduced into a lower portion of reaction zone 208 through vessel shell connection 58 located in a lower portion of reaction zone 208. A predominately liquid-phase feed catalyst liquid supply stream 3 is fed through conduit 240 and control valve 241 and is introduced into reaction zone 208 via feed inlet 242. Multi-phase reaction medium 11 comprises at least one liquid phase and a gas phase, as already disclosed, and is disposed predominantly in reaction zone 208. [00416] In FIG.3, distance “D” is the maximum horizontal inside diameter of reaction vessel shell 200, and distance “W” is the maximum horizontal diameter of reaction medium 11 within reaction vessel shell 200. Distance "H" is the height of reaction medium 11 having a horizontal diameter greater than W/3. Distance “L” is the inside height of reaction vessel shell 200 wherein the horizontal diameter is greater than D/3. Distance “Y” is the height of the gas- liquid disengaging zone 14 and of the disengaging section 202 of bubble column reactor vessel shell 200. [00417] In FIG.3, upper surface 210 of reaction medium 11 is defined along a horizontal plane that cuts through reaction vessel shell 200 at a vertical location where the contents of reaction vessel shell 200 transition from a gas- phase-continuous state to a liquid-phase-continuous state. Upper surface 210 is preferably positioned at the vertical location where the local time-averaged and horizontal area-averaged gas hold-up of a thin horizontal slice of the contents of reaction zone 208 is 0.9. [00418] In FIG.3, the purge gas feed stream 1 is flowing within conduit 51 and conforms to the disclosures made herein using FIG.1. Additional feed stream 2 is provided using multiple separate flows (i.e., olefin 2a, syngas mixture 2b, CO 2c, H2 2d, and inert gaseous compounds 2e), each independently determined and controlled according to disclosures made herein, that are flowing in individual conduits (i.e., olefin conduit 52a, syngas mixture conduit 52b, CO conduit 52c, H2 conduit 52d, and inert gaseous compound conduit 52e). [00419] In FIG.3, streams 1, 2a, 2b, 2c, 2d, and 2e are combined into a single flow stream 53 within conduit 54 and then conveyed into heating means 5 which is supplied with a heating fluid supply 6, e.g., steam, organic heat transfer medium, and so on, and heating fluid return 7. Heated feed stream 55 flows from heating means 5 via conduit 56 and enters bubble column reactor shell 200 via vessel shell connection 58, as already disclosed. [00420] The provision of heating means 5 is useful for heating reaction medium 11 to initiate reaction start-up. In addition, heating means 5 is often useful during periods when ambient conditions have cooler temperature, e.g., wintertime, particularly when olefin fraction of stream 53 is relatively small. In such conditions, ambient losses through insulation are greater; and cooling provided by the mass flow of stream 53 through reaction medium 11 exiting as spent gas stream 12 is often greater; and it may be that the heat of reactions within reaction medium 11 may not produce an overall cooling duty in the energy balance for reaction medium 11. In such conditions, heating means 5 is employed to increase the temperature of combined feed stream 56 in order that a small cooling duty is imposed upon reactor cooling means 20 in order to enable control of the temperature of reaction medium 11 at the preferred conditions already disclosed herein. Accordingly, it is preferred that stream 53 can be heated by heating means 5 to provide a temperature in stream 55 of at least about 80°C, more preferred at least about 110°C, still more preferred at least about 130°C, and most preferred at least 150°C. When the amount of aldehyde formed within reaction medium 11 is large enough to produce a significant cooling duty within reactor cooling means 20, then it is desirable to minimize or even stop the heating duty in heating means 5. In such cases, the temperature of combined feed streams 53 and 55 may be at or even less than ambient temperature. Optionally, heating means 5 may be located within reaction medium 11 inside of reactor shell 200, or heating means 5 may be provided as heat input transmitted through reactor shell 200, e.g., heat tracing and heat jacketing. [00421] Optionally, a feed distribution means 203 may be provided and disposed within the reaction vessel shell 200. Means 203 may be a perforated sieve plate, a perforated conduit sparger, or any other flow distributor means as known in the art. It is preferred that vessel shell connection 58 and, when provided, optional distribution means 203 both be located less than about 4*D, more preferred less than about 2*D, still more preferred less than about 1*D, and most preferred less than D/2 above the lowest portion of reaction vessel shell 200. When feed distribution means 203 is provided, it is preferred that the flowing pressure drop of the gas phase caused by feed distribution means is at least about 5 kPa, more preferred at least about 10 kPa, still more preferred at least about 20kPa, and most preferred at least 30 kPa (0.7, 1.5, 2.9, 4.4 psi); and it is preferred that the pressure drop is less than about 300 kPa, more preferred less than about 240 kPa, still more preferred less than about 180 kPa, and most preferred less than 120 kPa (44, 35, 26, 17 psi). However, when using the relatively small D and W and relatively large L/D and H/W ratio as is disclosed later herein, it is preferred to omit the mechanical complexity of distribution means 203 simply releasing feed stream 55 directly into reaction medium 11 via vessel shell connection 58. [00422] It is preferred to operate with a relatively large gas hold-up within reaction medium 11 in order to sustain greater mass transfer rates for dissolution of gaseous olefin, CO and H2 into the liquid phase while simultaneously reducing some adverse effects of having a greater mass of liquid phase, e.g., greater inventories of costly catalyst compounds, greater degradation rates of catalyst compounds, greater yield losses to heavy condensation products of aldehydes, greater capital cost, and so on. Accordingly, it is preferred that the time-average and volume averaged gas holdup of reaction medium 11 is at least about 2%, more preferred at least about 4%, still more preferred at least about 8%, and most preferred at least 16%. Other considerations comprising capital cost lead to a preferred gas hold- up within reaction medium 11 that is less than about 85%, more preferred less than about 75%, still more preferred less than about 65%, and most preferred less than 55%. [00423] Furthermore, it is preferred that the time-averaged gas-hold up within reaction medium 11 is in the ranges just disclosed when measured specifically at positions H/4, H/2, and 3*H/4 above the bottom of reaction medium 11, i.e., at the one-quarter height, one-half-height, and three-quarter heights within reaction medium 11. In cases where the mole fraction of olefin in feed stream 53 is relatively large and the mole fraction of inerts therein is relatively small, the total molar flow, and hence volumetric flow, of gas phase may be reduced considerably as portions of olefin, CO, and H2 are consumed to produce aldehyde while axially upwards from the bottom of reaction medium 11. In such cases, it is often preferred to keep the gas hold-up within preferred ranges by using a reduced cross-sectional diameter in at least a portion of the upper elevation of bubble column reactor shell 200. [00424] To obtain the disclosed preferred ranges of gas hold-up within reaction medium 11 when using compositions, pressures and temperatures disclosed herein, it is preferred that the superficial gas velocity within reaction medium 11 is at least about 0.01 meters per second, more preferred at least about 0.02 meters per second, still more preferred at least about 0.04 meters per second, and most preferred at least 0.06 meters per second; and it is also preferred that the superficial gas velocity is less than about 1.2 meters per second, more preferred less than about 0.9 meters per second, still more preferred less than about 0.6 meters per second, and most preferred less than 0.4 meters per second. The term "superficial gas velocity", as defined herein with reference to reaction medium 11, shall denote the volumetric flow rate of the gas phase of reaction medium 11 at an elevation in the reactor divided by the horizontal cross-sectional area of the reactor at that elevation. In cases where the mole fraction of olefin in feed stream 53 is relatively large and the mole fraction of inerts therein is relatively small, the total molar flow, and hence volumetric flow, of gas phase may be reduced considerably as aldehyde is produced rising axially from the bottom of the reactor. In such cases, it is often preferred to keep the superficial gas velocity within preferred ranges just disclosed by using a reduced cross-sectional diameter in at least a portion of the upper elevation of bubble column reactor shell 200. [00425] The preferred dimensions for D and W are at least about 0.075 meters, more preferred at least about 0.100 meters, still more preferred at least about 0.150 meters, and most preferred at least 0.200 m (3, 4, 6, 8 inches); and the preferred dimensions for D and W are less than about 6 meters, more preferred less than about 4 meters, still more preferred less than about 2 meters, and most preferred less than 1 meter (19.7, 13.1, 6.6, 3.1 feet). The diameter of a bubble column reactor inherently affects the superficial gas velocity, the gas hold-up, and the complex multi-phase mixing behaviors of the natural convection. [00426] In order to obtain a greater conversion of olefin into aldehyde product while operating with substantially one-pass mode of the gas phase as preferred and disclosed herein, the following ranges of height are preferred. It is preferred that height H of reaction medium 11 is at least about 4 meters, more preferred at least about 6 meters, still more preferred at least about 8 meters, and most preferred at least 10 meters. To limit adverse costs, e.g. greater capital, more production of unwanted heavy condensation products of aldehyde, more inventory of catalyst compounds, and so in, it is also preferred to limit height. It is preferred that height H of reaction medium 11 is less than about 75 meters, more preferred less than about 60 meters, still more preferred less than about 45 meters, and most preferred less than 30 meters. [00427] It is preferred that the height Y of gas-liquid disengaging zone 14 near the top of reactor vessel shell 200 is at least about 1 meters, more preferred at least about 2 meters, still more preferred at least about 2.5 meters, and most preferred at least 3 meters; and it is preferred that the height Y is less than about 12 meters, more preferred less than about 10 meters, still more preferred less than about 8 meters, and most preferred less than 6 meters. These can be operating upsets, bubble swarm eruptions, liquid phase foaming, and capital cost. [00428] It is preferred that the height L of reactor vessel shell 200 is at least about 6 meters, more preferred at least about 8 meters, still more preferred at least about 10 meters, and most preferred at least 12 meters; and it is preferred that height L is less than about 75 meters, more preferred less than about 60 meters, still more preferred less than about 45 meters, and most preferred less than 30 meters. These are in consideration of various combinations of preferred ranges of H and Y. [00429] The preferences for relatively tall reaction medium and vessel height and for relatively great superficial gas velocities and gas hold-ups coupled with what is often a relatively small flow of purge gas feed stream 1 and combined feed stream 53, when provided according to preferred composition ranges herein, lead to preferred aspect ratios for the bubble column reactor vessel shell and reaction medium. It is preferred that the ratio L/D is at least about 10, more preferred at least about 20, still more preferred at least about 25, and most preferred at least 30; and it is preferred that the ratio L/D is less than about 200, more preferred less than about 150, still more preferred less than about 125, and most preferred less than 100. It is preferred that the ratio H/W is at least about 8, more preferred at least about 17, still more preferred at least about 22, and most preferred at least 27; and it is preferred that the ratio H/W is less than about 187, more preferred less than about 139, still more preferred less than about115, and most preferred less than 90. [00430] The various preferences for H and superficial gas hold-up provide a desirably tall interfacial contact path and desirably great interfacial mixing between gas and liquid phases of the reaction medium. However, this necessarily expends pressure energy thereby reducing system pressure in upper portions of reaction medium 11 and reaction vessel 200. Accordingly, various combinations of H, Y, and superficial gas hold-up, coupled with the densities of gas and liquid phases, give rise to preferred ranges for pressure differential rising through reaction medium 11 and through reactor vessel shell 200. It is preferred that the pressure at the bottom of reaction medium 11 minus the pressure at reaction medium surface 210 and that the pressure at the bottom of reaction vessel shell 200 minus the pressure at the top of reactor vessel shell 200 are at least about 25 kPa, more preferred at least about 40 kPa, still more preferred at least about 55 kPa, and most preferred at least about 70 kPa (4, 6, 8, 10 psi). It is also preferred that the pressure differences are less than about 270 kPa, more preferred less than about 240 kPa, still more preferred less than about 210 kPa, and most preferred less than 180 kPa (39, 35, 31, 26 psi). [00431] For greater conversion of olefins passing upwards through reaction medium 11 in the bubble column reactor of FIG.3, it is preferred to limit the amount of back-mixing of gas phase from higher elevations to lower elevations, as previously disclosed herein. The gas phase of a bubble column reactor is not inherently in plug flow upwards. Although the relatively tall heights and large aspect ratios preferred herein for reaction medium 11 help to reduce desirably the amount of back-mixing of the gas phase as it travels upwards through reaction medium 11, the H/W ratio does not solely control the amount of axial back-mixing of the gas phase. The physical properties of the phases, the absolute values of W and of H independent of their ratio, and the superficial gas velocity also affect gas phase back-mixing. Thus, for some combinations of reaction medium composition, flows, and physical dimensions H and W, it is often preferred to provide optional gas RTD staging means 204 disposed within reaction medium 11 and attached to reaction vessel shell 200. As disclosed herein, such internals may be formed from an unlimited number of shapes and may be placed in any orientations and may be present in unlimited numbers. [00432] Horizontally disposed sieve trays are particularly preferred for gas RTD staging means 204. It is preferred that the trays have hole sizes in the range of at least about 1 millimeter, more preferred at least about 2 millimeters, still more preferred at least about 4 millimeters, and most preferred at least 6 millimeters. It is preferred that the trays have hole sizes in the range of less than about 200 millimeters, more preferred less than about 150 millimeters, still more preferred less than about 100 millimeters, and most preferred less than 50 millimeters. It is preferred that the holes on each tray occupy at least about 2%, more preferred at least about 4%, still more preferred at least about 8%, and most preferred at least 16% of the horizontal cross-sectional area of the reactor means 10 and reaction medium 11 at the elevation at which a particular tray is located. It is preferred that the holes on each tray occupy less than about 90%, more preferred less than about 70%, still more preferred less than about 50%, and most preferred less than 40% of the horizontal cross-sectional area of the reactor means 10 and reaction medium 11 at the elevation at which a particular tray is located. It is preferred that the holes are of uniform size in a single tray and are arranged in a regular pitch pattern approximately uniformly distributed across the horizontal area of the tray. However, all usage of multiple hole sizes with all possible distributions of holes within a single tray and usage of different sizes and distributions in different trays are disclosed by the inventors as aspects of the invention herein. [00433] Owing to the rising gas phase and the consumption of olefin, CO and H2 therein, there is necessarily developed an axial profile in various aspects of reaction medium 11, e.g. local pressure, local chemical compositions of the phases, and so on. In addition, convective mixing is developed within reaction medium 11 owing to the energy expended by the rising gas phase, and this contributes to the development of the axial profiles. Providing at least one gas RTD staging means 204 necessarily alters some or all of these axial profiles, notably including the local temperature of reaction medium 11 and of the gas and liquid phases thereof. [00434] In another aspect of the invention, it is preferred to provide a greater temperature in an upper portion of reaction medium in order to increase the concentration of aldehyde within spent gas stream 12 while providing a lesser temperature in a lower portion of reaction medium 11 in order to decrease degradation of catalyst compounds and formation of heavy condensation products the lower portion. This desirably increases the amount of aldehyde removed by a given flow of spent gas 12 operating in a substantially or entirely one-pass-gas-flow-self-stripping mode while also minimizing certain adverse effects of hotter temperatures in reaction medium 11. [00435] Accordingly, it is preferred that the time-averaged and volume- averaged temperature of a continuous portion of at least about 10 percent of reaction medium 11 located more than H/2 above the bottom of reaction medium 11 and reaction vessel shell 200 is hotter than the time-averaged and volume-averaged temperature of a continuous portion of at least about 10 percent of reaction medium 11 located less than H/2 above the bottom of reaction medium 11 and reaction vessel shell 200 by at least about 2°K, more preferred at least about 3°K, still more preferred at least about 4°K, and most preferred at least 5°K. However, it is also preferred to the temperature difference to less than about 30°K, more preferred less than about 25°K, still more preferred less than about 20°K, and most preferred less than 15°K. [00436] Provision of the temperature difference is accomplished balancing local aspects of the energy balance for reaction medium 11 and local aspects of the axial convective mixing. Aspects of the invention of particular importance comprise the olefin content and temperature of feed stream 55; the amount of aldehyde produced and the locus of the production within reaction medium 11, the disposition of cooling means 20 within reaction medium 11 and/or the vertical positioning of vessel shell connections 221 and 226; the interaction of H, W, superficial gas velocity, average gas hold-up in providing convective axial mixing of the gas and liquid phases of reaction medium 11; and the number, locations, and types of optional RTD staging means 204. Many of these aspects are set by chemistry and the compositions and temperatures of streams 2, 3 and 32 coupled with H and W of the reaction medium 11. Other aspects are controlled by the physical disposition of at least one cooling means 20 or at least one gas RTD staging means 204. Accordingly, it is preferred to locate at least one cooling means 20 below at least one gas RTD staging means 204, more preferred below at least 2 of the staging means, still more preferred below at least 3 of the staging means, and most preferred below at least 4 of staging means. In this embodiment, it is preferred to locate the lowest of the staging means 204 at least about H/8, more preferred at least about H/6, still more preferred at least about H/5, and most preferred at least H/4 above the bottom of reaction medium 11 and reaction vessel shell 200; and it is preferred to locate the lowest staging means less than about 7*H/8, more preferred less than about 3*H/4, still more preferred less than about 5*H/8, and most preferred less than H/2 above the bottom of reaction medium 11 reaction and reaction vessel shell 200. [00437] In another aspect of the invention, it is preferred to locate the introduction of at least a portion of streams 1 and 2 significantly above the bottom of reaction medium 11 and reactor vessel shell 200. Reasons for doing so comprise hydrodynamics (e.g., excessive local gas-hold up, control of axial mixing, and so on) and reaction kinetics (e.g. delaying dilution of olefin by non- olefin, delaying addition of a portion CO, higher concentration of which sometimes retard reaction rates, and so on). In a preferred embodiment, it is preferred to admit at least a portion of streams 2b, 2c, 2d and 2e at least about H/5, more preferred at least about H/4, still more preferred at least about H/3, and most preferred H/2 above the bottom of reaction medium 11 and reactor vessel shell 200. In another preferred embodiment, it is preferred to admit the preponderance or even all of inert gaseous compound stream 2e above these elevations. In another embodiment it is preferred to admit at least a portion of streams 2b, 2c, 2d, and 2e above at least one gas RTD staging means 204. In another preferred embodiment, it is preferred to admit the preponderance or even all of inert gaseous compound stream 2e above at least one gas RTD staging means 204. [00438] Optionally, a demister means 204 is provided disposed within disengagement zone 209 inside of disengaging section 202 of reaction vessel shell 200. It is preferred to locate demister means 204 above reaction medium surface 210 by at least about 250 millimeters, more preferred at least about 500 millimeters, still more preferred at least about 750 millimeters, and most preferred at least 1 meter. It is preferred that demister means 204 is located below vessel shell connection 205 and in a portion of the disengaging section 202 of vessel shell 200. It is preferred that demister means 204 is any physical impingement means for coalescing and removing aerosols, mists and droplets from a predominantly gaseous stream as is known in the art, e.g. layers of wire screens, layers of wire mesh, closely spaced impingement surfaces of various surface shapes and dispositions such as structured packing, random packing, chevron shapes, and so on. It is preferred that the outside of demister means 204 very closely matches the inside diameter of reaction vessel shell 200 present at the elevation which it is located, minimizing by-passing of gas flow around it. It is likely that an upper portion of reactor shell 200 has a maximum internal diameter less than D, e.g. a curved top head, but it is nonetheless preferred that demister means 204 has an outside diameter of at least about D/8, more preferred at least about D/4, still more preferred at least about D/2, and most preferred a diameter D. [00439] Bubble column reactor 200 includes gas outlet 205 located near or at the top of disengagement section 202 of reactor vessel shell 200 in an upper portion of disengagement zone 209 A predominantly gaseous effluent stream 12 is withdrawn via gas outlet 205 and conveyed through conduit 206 to a gas cooling and condensing means 31. [00440] Gas cooling and condensing means 31 comprises 2 separate heat exchange means 301 and 302. The heat exchange surfaces in heat exchange means 301 and 302 may be constructed in any manner as already disclosed in description of means 31 in FIG.1. Plate-coil and shell-and-tube heat exchangers designed as partial condensers and as dephlegmators are preferred, along with usage of suitable grades of stainless steel for heat exchange surfaces. Exchange means 301 and 302 are served by a warmer temperature cooling fluid supply 311 and return 312, e.g. cooling tower water, river water, and so on, and a colder temperature cooling fluid supply 313 and return 314, e.g. chilled water, chilled brine, chilled glycol, hydrofluorocarbon refrigerant fluids, and so on. [00441] Referring still to FIG. 3, gas and liquid are separated in a gravity separating means 303, as is known in the art. Vent gas flow 32 leaves via a conduit 321. Control valve 320 located in conduit 321is adjusted in response to a pressure measurement and control means 320a connected to disengaging section 202 or to conduit 206 in order to control the pressure within gas-liquid disengaging zone 14 within disengaging section 202. Liquid crude aldehyde product stream 33 leaves via a conduit 331. Control valve 332 located in conduit 331 is adjusted in response to a liquid level measurement and control means 332a connected with a lower portion of separator vessel 303 in order to control the liquid level therein. [00442] In one aspect of the invention, liquid reflux 340, conduit 341, and control valve 342 are not provided. The level of reaction medium 11 within reaction vessel shell 200 is allowed simply to seek its own steady state condition according to the temperature and composition of reaction medium 11, to the production of aldehyde and heavy condensation products of aldehydes within medium 11, and to the flow rate and composition of spent gas stream 12. Infrequent, relatively small flows of catalyst liquid supply 4 and catalyst liquid removal 3 are manually controlled, in accordance with frequency, duration and flow volume ratios disclosures herein, in order to maintain the level of reaction medium 11 in an acceptable range. [00443] In another aspect of the invention, liquid reflux 340, conduit 340, and control valve 342 are provided along with reaction medium level measurement and control means 342a. This allows more stable control of the level of reaction medium 11, which may be of use in favorably maximizing conversion of olefin to aldehyde; but flow of liquid reflux 340 also possibly reduces the maximum production rate of stream 33, and it also increases the concentration of aldehyde within reaction medium 11, which is likely to lead to an increased formation rate of heavy condensation products of aldehyde. [00444] In another aspect of the invention, different and perhaps complex control logic arrangements may be employed using control valves 225, 332, 342 plus adjustments in flows via streams 2, 3, and 4 in order to optimize the overall economic efficiency considering the flow rates and economic values of aldehyde product, undesirable byproducts, additional feed stream 2, and vent stream 32 plus the value of concentrating catalyst metal in liquid stream 4 compared to stream 3. [00445] When optional liquid reflux stream 340 is provided, it is preferred to locate optional vessel shell connection 343 in disengaging section 202 of reaction vessel shell 200. It is preferred to locate vessel shell connection 343 less than about 6 meters, more preferred less than about 5 meters, still more preferred less than about 4 meters, and most preferred less than 3 meters below the top of reaction vessel shell 200. It is preferred that reflux stream 340 is released into gas-liquid disengaging zone 14 at least about 500 millimeters, more preferred at least about 750 millimeters, still more preferred at least about 1,000 millimeters, and most preferred at least 1,500 millimeters below the top of disengaging zone 14 and the top of reaction vessel shell 200. When optional demister means 204 and optional liquid reflux stream 340is are both provided, it is preferred that optional liquid reflux stream 340 is released above or within demister means 204; and it is preferred to use a liquid gravity distribution means 344 to distribute stream 340 horizontally in a relatively uniform manner above or within demister means 204 according to apparatus and methods as is known in the art of scrubbers and distillation columns, e.g. multiple orifice conduits, sieve trays, baffled sieve trays, v-notch channels, and so on. [00446] FIG.3 discloses a preferred disposition for reactor cooling means 20 on the outside of bubble column reactor vessel shell 200. Advantages of this location comprise simplified and standardized mechanic construction and simplified periodic maintenance for reactor cooling means 20 plus avoiding usage of a tempered and/or non-corrosive heat transfer fluid for coolant streams 20 and 21. In FIG.3, it is preferred that coolant supply 21 and return 22 are a standard cooling water source such as cooling tower water, lake water, river water, and the like instead of a more costly coolant that is less sedimenting and less corrosive. [00447] For temperature control of reaction medium 11 located within reaction zone 208, the position of control valve 225, which is located in liquid return conduit 224, is adjusted in response to at least one temperature measurement and control means 225a, which is in contact with a portion of reaction medium 11 located within vessel section 201. When control valve 225 is at least partially open, multiphase reaction medium 11 flows by force of gravity into a gas-liquid separation means 220 via a vessel shell connection 221 located on vessel shell 200 in reaction zone 208and via conduit 222. Gas- liquid separation means 220 may have many different mechanical designs, and it is not necessary to completely separate all of the gaseous phase from the liquid portion of reaction medium 11. When separation means 220 removes at least a portion of gaseous mixture prevailing within reaction medium 11, a gravity-elevation-head pressure differential is set up in the liquid-rich portion of reaction medium 11b in conduit 223 and 11c in conduits 224 and 224a in comparison with the gravity-elevation-head in the corresponding, more aerated portion of reaction medium 11 within reaction zone 208. When control valve 225 is at least partly open, this gravity-elevation-head pressure differential causes a portion of liquid-rich reaction medium 11b to pass through cooling means 20 to produce a cooled portion of liquid-rich reaction medium 11c that is returned through conduits 224 and 224a into reaction vessel 200 via vessel shell connection 226. A more open position for valve 225 enables greater flow of a liquid-rich cooled reaction medium 11b and hence more cooling of reaction medium 11. [00448] In FIG.3, angle 221a is the included angle between vessel shell 200 and vessel shell connection 221. Angle 221a may be any angle between, but not including, 0° and 180° meaning that at least a portion of conduit 222 attached to vessel shell connection 221 may point in any direction from steeply downwards below horizontal ranging up to steeply upwards above horizontal. In a preferred embodiment, angle 221a is between 10° and 90° meaning that at least a portion of conduit 222 attached to vessel shell connection 221 is pointed horizontally or downwards. [00449] In a more preferred embodiment, angle 221a is about 90° and the majority, more preferably entirety, of conduit 222 is about horizontal. In this embodiment it is preferred to use optional return gas conduit 217 to vent a portion of gas-enriched reaction medium 11a up optional conduit 217 to re-enter reaction vessel 200 via vessel shell connection 218 located at a higher elevation than vessel shell connection 221. In this embodiment, it is preferred that vessel shell connection 218 be located within reaction zone 208 rather than in disengaging section 202 so that portion 11a does not derive a by-passing route upwards through optional conduit 217 for a portion of the gas phase of reaction medium 11 to reach spent gas stream 12. It is more preferred that vessel shell connection 218 is located below the top of reaction medium surface 210 by a distance at least about 1*D, still more preferred at least about 2*D, and most preferred at least 4*D. It is preferred that the location of vessel shell connection 218 is above vessel shell connection 221 by about 1*D. [00450] In a more preferred embodiment of FIG.3, angle 221a is between about 30° and 60° and conduit 221 is aligned linearly at the same angle. When the internal diameters of vessel shell connection 221 and conduit 222 are aptly sized, such a slope allows a gas-rich portion of reaction medium 11a and a liquid rich portion of reaction medium 11b to counterflow, more gas along and near the top of the conduit and more liquid along and near the bottom, time- averaged. In this embodiment, it is preferred that the diameters of vessel shell connection 221 and of conduit 222 are each such that the respective time- averaged superficial velocity of the liquid phase within each is less than about 0.50 meters per second, more preferred is less than about 0.30 meters per second, still more preferred less than about 0.20 meters per second, and most preferred less than 0.15 meters per second when providing the maximum required cooling duty in reactor cooling means 20. Such slow superficial velocities of the liquid phase enable substantial deaeration of reaction medium portion 11b and also control frictional flowing pressure drop within conduit 222. It is preferred that the superficial velocity of liquid phase in gas-liquid separation means 220, which may optionally be a linear extension of conduit 222, conform to the same preferred ranges of superficial liquid velocity. [00451] To balance flowing pressure drop and residence time of reaction medium portion 11b outside of reaction vessel shell 200, it is preferred that the velocities within conduits 223, 224, 224a, and 224b are less than about 2.5 meters per second, more preferred less than about 2.0 meters per second, still more preferred less than about 1.5 meter per second, and most preferred less than 1.3 meters per second; and it is preferred that the velocities are at least about 0.5 meters per second, more preferred at least about 0.6 meters per second, still more preferred at least about 0.7 meters per second, and most preferred at least 0.8 meters per second. [00452] It is preferred that the pressure drop of reaction medium 11b in reactor cooling means 20 be less than about kPa, more preferred is less than about kPa, still more preferred less than about kPa and most preferred less than kPa. In combination with the slow velocities disclosed for conduits 222, 223, 224, and 224a, this limits total pressure drop of reaction medium 11, 11a, and 11b in the external cooling means to that reasonably obtained by gravity driven liquid circulation, e.g., without the need for mechanical shaft work pumping to return cooled reaction medium 11c to the main portion of reaction medium inside reactor shell 200. [00453] It is preferred that vessel shell connection 221 be located above the bottom of reaction vessel shell 200 by at least about 2 meter, more preferred at least about 3 meters, still more preferred at least about 4 meters, and most preferred at least 6 meters. Greater elevations for vessel shell connection 221 increase the gravity-elevation-head pressure differential provided to the cooling liquid circulation flow, but care must be taken that upset operating conditions cannot lower the elevation of reaction medium surface below vessel shell connection 221. It is preferred that the distance between vessel shell connections 221 and 226 be at least about 2 meters, more preferred at least about 3 meters, still more preferred at least about 4 meters, and most preferred at least 5 meters. It is preferred that shell connection 226 be located at an elevation that is above the bottom of vessel shell 200 less than about 8*D, more preferred less than about 4*D, still more preferred less than about 2*D, and most preferred less than 1*D, for more of the heat of reaction and hence more of the cooling duty will be in lower portions of reaction medium 11 where the concentration of olefin is greater. [00454] Optionally, a portion or all of cooled medium 11c is returned to via optional conduit 224b. Optionally, the portion is mixed directly with feed stream 55 in conduit 56, or the portion is mixed with feed stream 55 in optional gas- liquid eductor 57. Both options provide additional differential pressure for the external cooling liquid circulation loop by consuming flow energy from stream 55; both options necessarily reduce somewhat the maximum pressure control setting available via means 320a, and both options provide a more highly aerated portion of cooled reaction medium 11d. Designs for optional eductor 57 are well known in the art and are commercially available. It is preferred that the attachment point 227, optionally to conduit 56 or to optional eductor 57, is below the bottom of reactor vessel shell 200 by at least about 0.25 meters, more preferred at least about 0.50 meters, still more preferred at least about 0.75 meters, and most preferred at least 1 meter. It is preferred that the attachment point 227 is below the bottom of reactor vessel shell 200 by less than about 8 meters, more preferred less than about 6 meters, still more preferred less than about 4 meters, and most preferred less than 3 meters. [00455] When cooling means 20 is disposed outside of reaction vessel shell 200, reaction medium portions 11a, 11b, and 11c are intentionally greatly de- aerated compared to portions of reaction medium 11 located inside of reaction zone 208 disposed within the shell. Often this adversely affects the selectivity of reactions, i.e., aldehyde isomer formation ratio and/or byproduct formation rate, and/or the decomposition rates of catalyst compounds, especially before passing through cooling means 20. Accordingly, it is preferred to design such an external reactor cooling means 20 such that the volume of all reaction medium outside of reaction vessel shell 200 is less than about 50%, more preferred less than about 40%, still more preferred less than about 30%, and most preferred less than 20% of reaction medium 11 within reaction vessel shell 200. Examples [00456] Examples 1-5 are calculated examples of a purge gas scrubbing reactor relating particularly to the continuous removal by gas-phase stripping only of an amount of aldehyde produced by liquid-phase hydroformylation plus the amount of said aldehyde that may have been present in the purge gas fed to the reactor system. Other aspects of the invention are in evidence as well. [00457] Examples 1-5 all employ a bubble column reaction system as shown in FIG.3. The reaction vessel has a nearly vertical, essentially cylindrical body with an inside diameter D of about 0.152 meters (6 inches). The height H of the bubble column reaction vessel is about 15.2 meters (50 feet) from Lower Tangent Line (LTL) to Upper TL (UTL). The vessel is fitted with about 2:1 elliptical heads at the top and bottom of the cylinder. Feed gas enters the reaction vessel via a nominal 0.05-meter diameter round conduit (nominal 2- inch pipe) oriented vertically at the lowest portion of the bottom head. Gas exits the reaction vessel via a nominal 0.05-meter diameter round conduit (2-inch nominal pipe) oriented vertically at the highest portion of the top head. [00458] The vessel is equipped with an external shell and tube exchanger placing the reactor liquid on the shell side and pressurized inhibited water circulation on the tube side. Reaction medium is drawn from the reaction vessel via a sloped conduit attached to the cylinder of the reaction vessel at an elevation of about 3 meters above the LTL. A portion of the gas flows upwards and back into the reaction vessel providing a denser, less-aerated 2-phase mixture that flows downwards via a conduit into and through the shell side of the exchanger oriented about horizontally and situated adjacent to the reaction vessel at an elevation about 2.5 meters above the LTL. The sizing of the exchanger contact area is large enough to remove the heat of hydroformylation provided by conversion of 1.5 kg-mole/hr of ethylene to propionaldehyde, providing a useful excess of heat exchange capability beyond the duties required in Example 1-5. The temperature of the water is adjusted through the range from about 40 to 120°C to provide start-up heating plus regulation of the continuous operating temperature of the reaction medium inside the reaction vessel as stated in the individual examples. The cooled, or heated, liquid flows from the exchanger via another conduit and enters back into the reaction vessel cylindrical shell at an elevation of about 0.15 meters above the LTL. [00459] The purge gas fed supplied into the purge gas scrubber reaction vessel and reaction medium is sourced from an associated primary hydroformylation reactor making propionaldehyde using a catalyst of Rhodium plus TCHP ligand. [00460] The composition of the purge gas feed varies in the 5 examples as shown in TABLE 1. Observe that for all cases the ethylene is at smaller concentration, and hence molar flow rate, than either the H2 and CO. Accordingly, the ethylene flow limits the amount of propionaldehyde that can be produced. [00461] It is also important that the purge gas feed contains a portion of propionaldehyde that has not been removed, e.g., propionaldehyde vapor that remains uncondensed when using cooling tower water coolant in a partial condenser operating on the recycle gas loop of the primary reactor. This propionaldehyde vapor within purge gas feed flow must be accounted when considering whether the purge gas reactor is self-stripping for propionaldehyde production by once through flow of the fed purge gas. [00462] The catalyst within the liquid phase of the reaction medium of the purge gas scrubber reactor for Examples 1-5 is Rh with TCHP ligand. The Rh in the purge gas reactor liquid is about 180 ppm by mass, and the TCHP ligand is present with a molar ratio of about 2:1 relative to the Rh. Importantly, segregation of the liquid phase of the purge gas scrubber reactor and the primary hydroformylation reactor enables independent optimization of the Rh and TCHP concentrations for the purge gas scrubber reactor duty in contrast to the primary reactor duty. For example, a larger concentration of Rh and a smaller molar excesses of TCHP can be used in the purge gas scrubber reactor to increase the ratio of hydroformylation reaction rate compared to the formation rate of heavy C6, C9, C12, etc. byproducts. Even with greater concentrations of Rh, the purge scrubber reactor has a vastly smaller mass of Rh than the main reactor. Even with reduced molar excess of TCHP, there is very low risk of Rh precipitation in the purge gas scrubber reactor because ligand poisons are highly unlikely to break through the primary reactor into purge gas feed. In addition, the segregation of the reactor liquids enables greater independence for optimizing the operating temperature of the purge gas scrubber reactor. [00463] The mass of reaction liquid within the reaction medium is about 67 kilograms. The reaction liquid comprises about 49 mole % propionaldehyde in all the examples. The reaction liquid comprises about 2 to 3 mole % of dissolved gases, varying among examples, when considering all reactant gases (ethylene, H2, CO) plus all of the more volatile and mostly inert compounds such as ethane, methane, nitrogen, argon, and carbon dioxide, i.e., the volatile, gaseous, inert compounds that accumulate in the recycle gas loop of the primary reactor gas and are the impetus for the purge gas flow. The balance of the reaction liquid is about 48 to 49 mole % and comprises the very small masses of catalyst compounds and mostly comprises the various heavy byproduct compounds derived from propionaldehyde by various aldehyde combination and decomposition reactions, e.g. products most typically with 3, 6, 9, 12, etc. carbon atoms. The number average molecular weight of the reaction liquid is about 130 to 140 atomic mass units for all the examples. [00464] The reaction vessel is fitted with 2 liquid conduits each with a nominal 0.025-m diameter (1-inch nominal pipe) and both attached within about 5D of the LTL. These conduits are used to supply liquid phase catalyst compounds prior to initial start-up, to make infrequent adjustments to the composition and chemical activity of the reaction medium, and to remove reaction liquid prior to maintenance activities. Both conduits are blocked by valves not allowing any liquid flows into or out of the reaction vessel during the substantially continuous operations described in Examples 1-5. [00465] In Examples 1-5, the vent gas exiting via the conduit atop the reaction vessel is passed through a heat removal exchanger to cool the top exit gas to condense most of the propionaldehyde into a crude liquid propionaldehyde stream. The temperature of the crude liquid propionaldehyde condensate and of the final vent gas leaving the partial condenser are both 40°C, obtained using cooling tower water as coolant. Optionally, a refrigerated cooling system could be used to condense additional, smaller amounts of crude propionaldehyde from the final vent gas before it is sent to boiler feed gas, flaring, or some other useful destination. [00466] In Examples 1-5, the amount of propionaldehyde carried from the top of the reaction medium and reaction vessel exceeds the sum of propionaldehyde produced by hydroformylation plus the amount of propionaldehyde contained in purge gas feed. That is, all 5 examples show operation where a once-through flow of purge gas provides more than enough gas-stripping carrying capacity to remove all produced propionaldehyde notwithstanding the nearly complete conversion of olefin reactant to propionaldehyde. [00467] Accordingly, a portion of the condensed crude propionaldehyde liquid is divided from the condensed crude propionaldehyde condensate and then returned by gravity to the reaction vessel and reaction medium to maintain the mass of liquid phase of the reaction medium. The net removal rate of product from the reaction vessel and medium is thereby matched to the production rate of product. The amount of crude propionaldehyde liquid returned to the reactor is adjusted by liquid flow valve via a control loop monitoring a differential pressure measurement between the LTL and UTL of the vessel. The mass of liquid in the vessel is directly related to this differential pressure, using the diameter of the vessel, virtually independent of the gas density and flow rate in the ranges described. [00468] Observe that although the mass of liquid in the reaction medium is maintained constant, the level (height) of the aerated, expanded reaction medium varies according to the gas flow rate and composition. A greater flow of purge gas feed of the same composition leads to a greater superficial gas velocity and hence greater gas hold-up at all elevations within the vessel, thereby causing the reaction medium to expand to a higher top elevation. A smaller content of ethylene gas in the same total molar flow rate cause a similar effect because a greater portion of CO and H2 remain unconsumed at the top of the vessel leading to a greater superficial velocity and medium expansion in upper portions of the vessel. Example 1 [00469] The bubble column reactor vessel of example 1 is internally open and unobstructed, i.e., the vessel does not include any horizontal baffling means such as sieve tray baffles. The conversion of ethylene to aldehyde is about 94.7% complete. The loss of unconverted ethylene from the scrubber reactor and into final purge vent gas has been reduced by a factor of about 20 compared to the purge gas fed from the primary reactor. [00470] The remaining gas flowing from the top of the reactor carries about 1.089 kg-mol/hr of propionaldehyde vapor from the vessel. This is in excess to the sum of the fed plus produced propionaldehyde by about 33%, and so the operation is self-stripping in a self-sustaining way without the need for recycle gas stripping or liquid withdrawal stripping. [00471] Because there is not a continuous flow of reaction liquid through the vessel, the composition of catalyst compounds is uniform axially. Furthermore, the bubble column reactor has a strong axial mixing current relative to the total heat of reaction and most of the propionaldehyde formation and most of the heat of reaction are in the lower half of the reaction medium where the cooling by incoming feed gas and liquid exchanger surface are located, and the reaction medium is effectively isothermal axially. [00472] The only significant axial concentration gradients are for the 3 fed reactant gases, in both gas and liquid phases. The gas phase CO is reduced from 25.0 mole % to 21.2 mole %, H2 from 42.0 mole % to 39.2 mole %, and ethylene from 5.50 mole % to 0.31 mole %. [00473] The top of the reaction medium is expanded to an elevation of about 6.3 meters. Owing to the tall, very skinny H:D of the reaction medium, the reactor is converting ethylene with the equivalent of about 3 to 4 equal subdivisions of the reaction medium arranged in series on the gas phase reactants only, despite being physically unbaffled and effectively without staging on liquid phase temperature, liquid phase catalyst concentration, liquid phase aldehyde concentration, liquid phase concentration of C6, C9, etc. heavy byproduct, and total dissolved gases in the liquid, which are dominated by light inerts far more than the dissolved CO, H2 and ethylene. Example 2 [00474] The bubble column reactor vessel is the same as Example 1 with the addition of 8 sieve tray horizontal baffles at elevations of about 2.3 m, 2.8 m, 3.3 m, 3.8 m, … 5.8 m above the LTL. This provides additional staging of the gas phase reactants flowing upwards. However, because so little ethylene is remaining and so little propionaldehyde is formed above 2.8 m in the reaction medium and because this limited reaction releases so little additional heat above 2.8 m, the flow restriction that the sieve trays impose on the liquid phase does not result in exacerbated axial gradients of liquid composition and temperature, other than the axial gradients in dissolved reactant gases, chiefly for the ethylene. [00475] The propionaldehyde production is increased compared to Example 1 up to about 0.463 kg-mole/hr providing an ethylene conversion of about 97.5%. The loss of unconverted ethylene has been reduced about in half compared to Example 1. [00476] Although slightly more propionaldehyde production is now removed with slightly less amounts of ethylene, CO and H2, remaining in the reactor overhead outlet flow, the self-stripping ratio is still a very comfortable 130% of fed plus produced propionaldehyde. The purge gas scrubber system retains enough excess stripping capacity to adjust to typical operational instabilities in the composition of purge gas feed without accumulation of unstripped liquid propionaldehyde in the purge scrubber reactor; and condensed crude propionaldehyde liquid is refluxed by gravity into the reaction vessel to maintain the mass of the reaction medium. Example 3 [00477] The bubble column reactor vessel and system is the same as Example 2, complete with 8 sieve tray gas staging baffles in the upper elevations as noted. However, the mole fraction of ethylene in the feed gas is increased to 7.0 mole %. [00478] Table 1 shows that this causes a greater production of aldehyde and a small sag in ethylene conversion to 97.2% in a single pass of the gas phase upwards through the reaction medium. [00479] The combination of a greater aldehyde amount leaving in the reactor overhead gas flow plus reduced quantities of ethylene, CO and H2 for stripping gas leads to a narrower balance on the excess stripping capacity using once- through gas flow. The propionaldehyde stripping balance is still adequate for normal operations, with stripping equal to 107% of fed plus produced propionaldehyde, but further increases in ethylene feed fraction or increasing the condensing temperature of the purge gas feed from the primary reactor, e.g., due to exchanger fouling or hotter cooling water, can place the stripping balance in a bind. Example 4 [00480] The bubble column reactor vessel and system is the same as Example 2 but with a further increase of ethylene to 8.0 mole % in the purge gas fed to the scrubber reactor plus the addition of about 0.680 kg-mole/hr of methane comingled with the purge gas feed, for reasons discussed below. [00481] The propionaldehyde production is about 0.650 kg-mol/hr. The ethylene conversion sags a bit more down to about 94.2%, placing it slightly below the conversion obtained on Example 1 with only 5.5 mole % ethylene fed to the reactor without gas staging sieve tray baffles. This conversion sag is partly due to the increased demand for propionaldehyde production, analogous to the shift from Example 2 to Example 3; and it is also partly due to the reactant dilution caused by the addition of methane, i.e., the methane addition dilutes the concentration and partial pressures of reactant gases ethylene, CO and H2 at all elevations within the reaction vessel and medium. [00482] The utility provided by the added methane comingled with the purge gas feed is that the methane exits overhead to provide additional gas stripping capacity and without reducing the temperature at the top of the reaction vessel and medium. With the added methane, the self-stripping ratio is increased to 109% even with 8 mole % ethylene in the purge feed gas. [00483] As disclosed in the invention, any non-reacting volatile gas can be added to purge gas feed to supplement the once-through gas stripping capacity, avoiding the need for a gas recycle compressor or liquid stripping operation. However, methane is particularly suitable since the final vent purge gas from the overhead condenser often goes to a boiler to recover fuel value. [00484] The diluting effect that methane has on reactivity up the entire axial height can be reduced by adding the methane into the top of the reactor rather than comingling with purge gas feed into the reactor bottom. For example, the methane supplement can be added between sieve tray 7 and 8 counting upwards from the LTL. The disadvantage of this elevated addition point for the “dry” methane gas is the sensible and latent cooling effect that is produced, which in turn reduces the mole fraction of propionaldehyde vapor in the gas phase, at an axial position relatively isolated from significant heat of reaction. This drawback of elevated addition of the methane can be corrected by providing an additional heat exchanger to preheat the fed methane, even to temperatures hotter than the liquid phase of the reaction medium. Example 5 [00485] Example 5 shows a different way to cope with an ethylene concentration of about 8 mole % in the purge gas feed while retaining gas flow once through and gas-only stripping of product. No supplementary methane stripping gas is added. Instead, the temperature of the reaction medium is increased from about 107 to about 115°C. The combination of less dilution of reactants by methane plus increased hydroformylation reaction kinetics caused by hotter temperature increases the formation of propionaldehyde to about 0.678 kg-mole/hr. The greater formation of propionaldehyde increases the required gas stripping rate, while at the same time further reducing the remaining ethylene, CO and H2 components of overhead stripping gas. [00486] However, the increase in vapor pressure of propionaldehyde at the hotter temperature is more than enough to compensate, and the once-through gas-only self-stripping ratio is 115%. [00487] An adverse effect of the hotter reaction temperature is increased formation rates for C6, C9, etc. heavy byproducts. Over a period of days, this increased formation rate tends toward additional accumulation of heavy byproduct mass in the reaction medium. However, increasing temperature also raises the vapor pressures of the heavy byproducts and this increase is on a steeper slope with temperature than is the vapor pressure of propionaldehyde. This helps offset the accumulation of heavy liquids by removing heavy byproducts at a greater ratio to propionaldehyde production. In addition, the formation of heavy byproducts is about second order in the liquid phase composition of propionaldehyde and first order in the total mass of liquid phase. The formation and removal balance for heavy byproducts is small enough at 115°C that small reductions in the propionaldehyde liquid concentration below 49 mole % and/or reductions in the mass of liquid phase within the reaction vessel are enough to bring the formation and stripping of heavy byproducts into balance without reducing the propionaldehyde stripping ratio below 100%.

Claims

CLAIMS: 1. A hydroformylation composition produced by the process comprising: (a) forming a multiphase reaction medium comprising a liquid phase and a gas phase in a hydroformylation reactor, wherein the multiphase reaction medium comprises one or more olefins, molecular hydrogen (H2), carbon monoxide (CO), and a catalyst; (b) reacting at least a portion of the one or more olefins, the H2, and the CO in the multiphase reaction medium to thereby form one or more aldehydes via hydroformylation; in which the partial pressure of CO in the reactor outlet gas stream divided by the partial pressure of CO in the reactor gas inlet ranges from 0.03 to 0.85. (c) withdrawing from the hydroformylation reactor at least a portion of a gas phase effluent comprising at least a portion of the one or more aldehydes and non-aldehyde components; and the ratio of molar flow rate of said non-aldehyde gas phase effluent components to the molar formation rate of said one or more aldehydes is at least 14:1 and not more than 36:1. (d) cooling at least a portion of the gas phase effluent to form at least a portion of cooled gas phase effluent and at least a portion of condensed crude aldehyde liquid phase comprising at least a portion of at least one aldehyde and at least a portion of at least one olefin dissolved therein; (e) separating and compressing at least a portion of the cooled gas phase effluent to form a gas phase recycle gas; (f) returning at least a portion of the recycle gas of step (e) to the multiphase reaction medium of step (b); (g) returning at least a portion of condensed crude aldehyde liquid to the hydroformylation reactor of step (a); (h) contacting in a syngas scrubber at least a portion of a gas phase raw syngas feed comprising H2, CO, and at least one syngas impurity with at least a portion of the condensed crude aldehyde liquid of step (d) to form a liquid phase of improved crude aldehyde and a gas phase of improved syngas feed wherein at least a portion of at least one of the dissolved olefins of the condensed crude aldehyde liquid is transferred into the improved syngas feed; (i) supplying at least a portion of the improved syngas feed from step (g) to the multiphase reaction medium of step (b) or to the recycle gas of step (e); (j) supplying at least a portion of at least one olefin feed supply to at least a portion of the reaction medium of step (b) or to at least a portion of the recycle gas of step (e) or to syngas-recycle gas mix or to improved syngas; (k) optionally, separating at least a portion of cooled gas phase effluent of step (d) or at least a portion of the recycle gas phase effluent of step (e) to form a purge gas; wherein at least part of the process is completed in previously used equipment.

2. The hydroformylation composition produced by the process according to Claim 1 wherein the molar ratio of said non-aldehyde components to said one or more aldehydes in said gas phase effluent is at least 15:1 and/or not more than 35:1.

3. The hydroformylation composition produced by the process according to Claims 1-2 wherein the molar ratio of said non-aldehyde components to said one or more aldehydes in said gas phase effluent is at least 16:1 and/or not more than 35:1.

4. The hydroformylation composition produced by the process according to Claims 1-3 wherein the molar ratio of said non-aldehyde components to said one or more aldehydes in said gas phase effluent is at least 17:1 and/or not more than 30:1.

5. The hydroformylation composition produced by the process according to Claims 1-4 wherein the molar ratio of said non-aldehyde components to said one or more aldehydes in said gas phase effluent is at least 18:1 and/or not more than 35:1.

6. The hydroformylation composition produced by the process according to Claims 1-5 wherein the molar ratio of said non-aldehyde components to said one or more aldehydes in said gas phase effluent is at least 20:1 and/or not more than 25:1.

7. The hydroformylation composition produced by the process comprising: (a) carrying out at least one hydroformylation reaction in a main reactor to thereby form a first quantity of one or more aldehydes; (b) withdrawing an olefin-containing gaseous effluent from said main reactor; (c) contacting at least a portion of said olefin-containing gaseous effluent with a catalyst-containing liquid phase medium in a purge gas reactor under conditions sufficient to form, via hydroformylation, a second quantity of one or more aldehydes in said liquid phase medium; and (d) stripping at least a portion of said aldehydes out of said liquid phase medium produced in step c) using a gas phase stripping medium to thereby produce a gaseous purge reactor effluent comprising at least a portion of said one or more aldehydes stripped out of said liquid phase medium, wherein less than 25 (10, 5, 1, 0) mole percent of said gaseous purge reactor effluent is recirculated back to said scrubbing reactor, wherein said one or more aldehydes are stripped out of said liquid phase medium in said stripping of step (d) at an average molar stripping rate that is at least 80 percent of the average molar rate of formation of said one or more aldehydes in step (c); wherein said purge gas comprises one or more olefins, carbon monoxide (CO), and molecular hydrogen (H2) in respective amounts such that the molar ratio of combined CO plus H2 to said one of more olefins is at least 3:1.

8. The hydroformylation composition produced by the process according to Claim 7 wherein wherein less than 10 mole percent of said gaseous purge reactor effluent is recirculated back to said scrubbing reactor.

9. The hydroformylation composition produced by the process according to Claims 7-8 wherein wherein less than 5 mole percent of said gaseous purge reactor effluent is recirculated back to said scrubbing reactor.

10. The hydroformylation composition produced by the process according to Claims 7-9 wherein wherein less than 1 mole percent of said gaseous purge reactor effluent is recirculated back to said scrubbing reactor.

11. The hydroformylation composition produced by the process according to Claims 7-10 wherein said purge gas comprises one or more olefins, carbon monoxide (CO), and molecular hydrogen (H2) in respective amounts such that the molar ratio of combined CO plus H2 to said one of more olefins is at least 4:1.

12. The hydroformylation composition produced by the process according to Claims 7-11 wherein said purge gas comprises one or more olefins, carbon monoxide (CO), and molecular hydrogen (H2) in respective amounts such that the molar ratio of combined CO plus H2 to said one of more olefins is at least 5:1.

13. The hydroformylation composition produced by the process according to Claims 7-12 wherein said purge gas comprises one or more olefins, carbon monoxide (CO), and molecular hydrogen (H2) in respective amounts such that the molar ratio of combined CO plus H2 to said one of more olefins is at least 6:1.

14. The hydroformylation composition produced by the process according to Claims 7-13 wherein said one or more aldehydes are stripped out of said liquid phase medium in said stripping of step (d) at an average molar stripping rate that is at least 90 percent of the average molar rate of formation of said one or more aldehydes in step (c).

15. The hydroformylation composition produced by the process according to Claims 7-14 wherein said one or more aldehydes are stripped out of said liquid phase medium in said stripping of step (d) at an average molar stripping rate that is at least 95 percent of the average molar rate of formation of said one or more aldehydes in step (c).